Method for converting a hydrocarbonaceous material to a fluid hydrocarbon product comprising p-xylene

ABSTRACT

The invention relates to a method for producing a fluid hydrocarbon product, and more specifically, to a method for producing a fluid hydrocarbon product via catalytic pyrolysis. The reactants comprise hydrocarbonaceous materials (e.g., biomass). The catalyst comprises a zeolite catalyst treated with a silicone compound. The product comprises p-xylene.

This application claims priority under 35 U.S.C. §119(e) to U.S.Provisional Application Ser. No. 61/655,605 filed Jun. 5, 2012. Thedisclosure in this provisional application is incorporated herein byreference.

STATEMENT REGARDING FEDERALLY SPONSORED RESEARCH OR DEVELOPMENT

The U.S. Government has a paid-up license in this invention and theright in limited circumstances to require the patent owner to licenseothers on reasonable terms as provided for by the terms of Grant No.CBET-0747996 awarded by the National Science Foundation.

FIELD OF INVENTION

This invention relates to a method for converting a hydrocarbonaceousmaterial to a fluid hydrocarbon product comprising p-xylene, and morespecifically, to a method for converting a hydrocarbonaceous material toa fluid hydrocarbon product comprising p-xylene via catalytic pyrolysis.

BACKGROUND

p-Xylene is used as a starting material for plasticizers and polyesterfibers. The oxidation of p-xylene is used to commercially synthesizeterephthalic acid. Further esterification of the acid with methanolforms dimethyl terephthalate. Both monomers may be used in theproduction of polyethylene terephthalate (PET) plastic bottles andpolyester clothing.

SUMMARY

Catalytic pyrolysis, including catalytic fast pyrolysis (CFP), is aprocess that may be used to convert a hydrocarbonaceous material (e.g.,biomass) into a fluid hydrocarbon product using rapid heating rates inthe presence of a catalyst. With the inventive method, the fluidhydrocarbon product comprises p-xylene, and may further compriseadditional aromatics, olefins, and the like.

With this invention, a hydrocarbonaceous material may be fed to areactor (e.g., a fluidized-bed reactor) where the hydrocarbonaceousmaterial may first thermally decompose to form one or more pyrolysisproducts. The pyrolysis products may comprise one or more pyrolysisvapors. These pyrolysis products may react in the presence of a modifiedzeolite catalyst to form one or more aromatic compounds as well asolefin compounds, CO and CO₂. The modified zeolite catalyst may comprisepores with pore mouth openings that have been reduced in size. Thepyrolysis products may enter the pores in the modified zeolite catalystwhere they may undergo reaction. The products formed in the catalystpores may then diffuse out of the pores. The aromatic compounds maycomprise p-xylene or xylenes with a relatively high selectivity towardsp-xylene. Advantages of this process include: 1) all the desiredchemistry may occur in a single-step process, 2) the process may use arelatively inexpensive zeolite catalyst, and 3) p-xylene may be producedat a relatively high level of production.

p-Xylene may be the most valuable of the xylenes (i.e., o-, m- andp-xylenes). However, during the catalytic pyrolysis of varioushydrocarbonaceous materials, the xylenes may be formed with the m-xyleneselectivity and/or o-xylene selectivity being the same as or higher thanthe p-xylene selectivity. The p-xylene that is produced may alsoisomerize to m-xylene and/or o-xylene. As a result, xylenes withundesirably high selectivities to m-xylene and/or o-xylene may beformed. The problem therefore is to provide a method that allows for theproduction of p-xylene or xylenes with a relatively high selectivity top-xylene. This invention provides a solution to this problem. Thepresent invention provides for the use of a zeolite catalyst that hasbeen modified to improve selectivity towards p-xylene. The zeolitecatalyst may comprise catalytic sites on the external surface of thecatalyst and pores with pore mouth openings. The pores may containinternal catalytic sites, some of which may be positioned near the poremouth openings, and some of which may be internal catalytic sitespositioned away from the pore mouth openings. The catalyst may bemodified by treating it with an effective amount of a silicone compoundto reduce the size of the pore mouth openings and to render at leastsome of the catalytic sites on the external surfaces of the catalystinaccessible to the reactants. The treatment process may also be used torender at least some of the catalytic sites in the pores near the poremouth openings inaccessible to the reactants.

The zeolite catalyst may be treated by applying a silicone compound tothe surface of the catalyst. Application of the silicone compound mayreduce the size of the pore mouth openings in the catalyst as well ascover or obscure catalytic sites on the external surface of the catalystas well as inside the pores of the catalyst near the pore mouthopenings. The covering of the catalytic sites with the treatment layermay inhibit and/or extinguish their catalytic activity. Catalytic siteson the external surface of the catalyst as well as catalytic sites inthe pores near the pore mouth openings are believed to be unselective inthe production of xylenes, while the catalytic sites in the pores awayfrom the pore mouth openings appear to provide for an increase in paraselectivity. Inhibition or extinction of the activity of the catalyticsites on the external surface of the catalyst and in the pores of thecatalyst near the pore mouth openings may increase the proportion of thecatalytic reaction that occurs within the zeolite pores away from thepore mouth openings, hence an increase in selectivity to p-xylene.

The pores with pore mouth openings that have been reduced in size mayallow for an increase in para selectivity for the xylenes. This may bedue to the fact that the reduced pore mouth openings may allow p-xyleneto diffuse out of the pores while the diffusion of m-xylene and o-xylenemay be restricted. This is illustrated in FIG. 2. In FIG. 2, a zeolitecatalyst is shown that has a pore mouth opening that has been reduced insize to allow p-xylene, but not m- or o-xylene, to diffuse out of thezeolite pore.

This invention relates to a method for producing a fluid hydrocarbonproduct comprising p-xylene from a hydrocarbonaceous material,comprising: feeding the hydrocarbonaceous material to a reactor;pyrolyzing within the reactor at least a portion of thehydrocarbonaceous material under reaction conditions sufficient toproduce a pyrolysis product; and catalytically reacting at least aportion of the pyrolysis product under reaction conditions in thepresence of a zeolite catalyst to produce the fluid hydrocarbon product;the zeolite catalyst comprising pores with pore mouth openings andcatalytic sites on the external surface of the catalyst, and aneffective amount of a treatment layer derived from a silicone compoundto reduce the size of the pore mouth openings and to render at leastsome of the catalytic sites on the external surface of the catalystinaccessible to the pyrolysis product. Catalytic sites may be positionedin the pores near the pore mouth openings and the treatment layer mayrender at least some of these catalytic sites inaccessible to thepyrolysis product.

With the present invention, the selectivity to p-xylene in the xylenesmay be up to 100% when the only active catalytic sites are within thecatalyst pores, and the only xylene to diffuse from the catalyst poresis p-xylene. In an embodiment, p-xylene may be produced in preference too-xylene and/or m-xylene, but some o-xylene and/or m-xylene maynevertheless be produced.

With the present invention, the fluid hydrocarbon product produced usingthe foregoing method may comprise xylenes and may be characterized by ap-xylene selectivity in the xylenes of at least about 40%, or at leastabout 45%, or at least about 50%, or at least about 55%, or at leastabout 60%, or at least about 65%, or at least about 70%, or at leastabout 75%, or at least about 80%, or at least about 85%, or at leastabout 90%.

The zeolite catalyst may be treated with the silicone compound to reducethe size of the pore openings, and cover or obscure catalytic sites onthe external surface of the catalyst. This treatment process may also beused to cover or obscure catalytic sites in the pores near the poremouths openings. The catalytic sites may also be referred to as acidsites. The covered or obscured catalytic sites may be referred to asdeactivated catalytic sites. The silicone compound may have a molecularsize that is incapable of entering the pores of the catalyst. During thecatalyst treatment process, the silicone compound may be applied to thecatalyst and subsequently calcined. This process may be repeated untilthe desired level of treatment is provided. The fraction of catalyticsites on the external surface of the catalyst that may be deactivated bytreatment with the silicone compound may be at least about 15%, or atleast about 25%, or at least about 35%, or at least about 45%, or atleast about 55%, or at least about 65%, or at least about 75%, or atleast about 85%, or at least about 90%, or at least about 95%, or atleast about 98%, or at least about 99%, of the available catalyticsites. The fraction of catalytic sites in the pores of the catalyst nearthe pore mouth openings that may be deactivated by treatment with thesilicone compound may be at least about 20%, or at least about 25%, orat least about 30%, or at least about 33%, or at least about 35%, or atleast about 40%, or at least about 45%, or at least about 49%, or atleast about 50%, or at least about 55%, or at least about 60%, or atleast about 65%, or at least about 70%, or at least about 75%, or atleast about 80%, or at least about 85%, or at least about 90%, or atleast about 95%, or at least about 99% of the available catalytic sites.The average diameter of the pore mouth openings subsequent to treatmentwith the silicone compound may be in the range from about 5 to about 10angstroms, or from about 5.2 to about 7.4 angstroms. The zeolitecatalyst may comprise silica and alumina. The silica to alumina molarratio may be in the range from about 10:1 to about 50:1, or in the rangefrom about 10:1 to about 40:1, or in the range from about 10:1 to about20:1, or about 15:1. The zeolite catalyst may further comprise nickel,platinum, vanadium, palladium, manganese, cobalt, zinc, copper,chromium, gallium, an oxide of one or more thereof, or a mixture of twoor more thereof.

The silicone compound may comprise a compound containing at least onegroup represented by the formula

The silicone compound may be represented by the formula:

wherein R₁ and R₂ independently comprise hydrogen, halogen, hydroxyl,alkyl, alkoxyl, halogenated alkyl, aryl, halogenated aryl, aralkyl,halogenated aralkyl, alkaryl or halogenated alkaryl; and n is a numberthat is at least 2. R₁ and/or R₂ may comprise methyl, ethyl or phenyl. nmay be a number in the range from about 3 to about 1000.

The silicone compound may have a number average molecular weight in therange from about 80 to about 20,000, or from about 150 to 10,000.

The silicone compound may comprise dimethylsilicone, diethylsilicone,phenylmethylsilicone, methylhydrogensilicone, ethylhydrogen silicone,phenylhydrogen silicone, methylethyl silicone, phenylethyl silicone,diphenyl silicone, methyltrifluoropropyl silicone, ethyltrifluoropropylsilicone, polydimethyl silicone, tetrachloro-phenylmethyl silicone,tetrachlorophenylethyl silicone, tetrachlorophenylhydrogen silicone,tetrachlorophenylphenyl silicone, methylvinyl silicone, hexamethylcyclotrisiloxane, octamethyl cyclotetrasiloxane, hexaphenylcyclotrisiloxane, octaphenyl cyclotetrasiloxane, or a mixture of two ormore thereof.

The silicone compound may comprise a tetraorthosilicate. The siliconecompound may comprise tetramethylorthosilicate, tetraethylorthosilicate,or a mixture thereof.

The reactor may comprise a continuously stirred tank reactor, a batchreactor, a semi-batch reactor, a fixed bed reactor or a fluidized bedreactor. Fluidized bed reactors may be particularly advantageous.

The hydrocarbonaceous material may comprise a solid hydrocarbonaceousmaterial, a semi-solid hydrocarbonaceous material, a liquidhydrocarbonaceous material, or a mixture of two or more thereof. Thehydrocarbonaceous material may comprise biomass. The hydrocarbonaceousmaterial may comprise plastic waste, recycled plastics, agriculturalsolid waste, municipal solid waste, food waste, animal waste,carbohydrates, lignocellulosic materials, xylitol, glucose, cellobiose,hemi-cellulose, lignin, sugar cane bagasse, glucose, wood, corn stover,or a mixture of two or more thereof. The hydrocarbonaceous material maycomprise furan and/or 2-methylfuran. The hydrocarbonaceous material maycomprise pinewood. The hydrocarbonaceous material may comprise pyrolysisoil derived from biomass, a carbohydrate derived from biomass, analcohol derived from biomass, a biomass extract, a pretreated biomass, adigested biomass product, or a mixture of two or more thereof. Mixturesof two or more of any of the foregoing may be used as thehydrocarbonaceous feed material.

The reactor may be operated at a temperature in the range from about400° C. to about 600° C., or from about 425° C. to about 500° C., orfrom about 440° C. to about 460° C.

The hydrocarbonaceous material may be fed to the reactor at a massnormalized space velocity of up to about 3 hour⁻¹, or up to about 2hour⁻¹, or up to about 1.5 hour⁻¹, or up to about 0.9 hour⁻¹, or in therange from about 0.01 hour⁻¹ to about 3 hour⁻¹, or in the range fromabout 0.01 to about 2 hour⁻¹, or in the range from about 0.01 to about1.5 hour⁻¹, or in the range from about 0.01 to about 0.9 hour⁻¹, or inthe range from about 0.01 hour⁻¹ to about 0.5 hour⁻¹, or in the rangefrom about 0.1 hour⁻¹ to about 0.9 hour⁻¹, or in the range from about0.1 hour⁻¹ to about 0.5 hour⁻¹.

The reactor may be operated at a pressure of at least about 100 kPa, orat least about 200 kPa, or at least about 300 kPa, or at least about 400kPa. The reactor may be operated at a pressure below about 600 kPa, orbelow about 400 kPa, or below about 200 kPa. The reactor may be operatedat a pressure in the range from about 100 to about 600 kPa, or in therange from about 100 to about 400 kPa, or in the range from about 100 toabout 200 kPa.

The method may be conducted under reaction conditions that minimize cokeproduction. The pyrolysis product may be formed with less than about 30wt %, or less than about 25 wt %, or less than about 20 wt %, or lessthan about 15 wt %, or less than about 10 wt %, of the pyrolysis productbeing coke.

The method may further comprise the step of recovering the fluidhydrocarbon product. The fluid hydrocarbon product may further comprise,in addition to p-xylene, other aromatic compounds and/or olefincompounds. The fluid hydrocarbon product may further comprise benzene,toluene, ethylbenzene, methylethylbenzene, trimethylbenzene, o-xylene,m-xylene, indanes naphthalene, methylnaphthelene, dimethylnaphthalene,ethylnaphthalene, hydrindene, methylhydrindene, dimethylhydrindene, or amixture of two or more thereof.

The carbon yield of aromatics in the fluid hydrocarbon product may be atleast about 10%, or at least about 15%, or at least about 22%. Thecarbon yield of olefins in the fluid hydrocarbon product may be at leastabout 3%, or at least about 6%, or at least about 9%, or at least about12%. The mass yield of p-xylene may be at least about 1.5 wt %, or atleast about 2 wt %, or at least about 2.5 wt %, or at least about 3 wt%.

The catalytically reacting step may comprise a dehydration,decarbonylation, decarboxylation, isomerization, oligomerization and/ordehydrogenation reaction.

The pyrolyzing step and the catalytically reacting steps may be carriedout in a single vessel. Alternatively, the pyrolyzing step and thecatalytically reacting steps may be carried out in separate vessels.

Other advantages and novel features of the present invention will becomeapparent from the following detailed description of various non-limitingembodiments of the invention when considered in conjunction with theaccompanying figures. In cases where the present specification and adocument incorporated by reference include conflicting and/orinconsistent disclosure, the present specification shall control.

BRIEF DESCRIPTION OF THE DRAWINGS

Non-limiting embodiments of this invention will be described by way ofexample with reference to the accompanying figures, which are schematicand are not intended to be drawn to scale. In the figures, eachidentical or nearly identical component illustrated is typicallyrepresented by a single numeral. For purposes of clarity, not everycomponent is labeled in every figure, nor is every component of eachembodiment of the invention shown where illustration is not necessary toallow those of ordinary skill in the art to understand the invention. Inthe figures:

FIGS. 1A and 1B are schematic illustrations of a CFP process forconverting a solid hydrocarbonaceous material to a fluid hydrocarbonproduct.

FIG. 2 is a schematic illustration showing a pore with a pore mouthopening that has been reduced in size and the diffusion of p-xylene outof the pore.

DETAILED DESCRIPTION

All ranges and ratio limits disclosed in the specification and claimsmay be combined in any manner. It is to be understood that unlessspecifically stated otherwise, references to “a,” “an,” and/or “the” mayinclude one or more than one, and that reference to an item in thesingular may also include the item in the plural.

The phrase “and/or” should be understood to mean “either or both” of theelements so conjoined, i.e., elements that are conjunctively present insome cases and disjunctively present in other cases. Other elements mayoptionally be present other than the elements specifically identified bythe “and/or” clause, whether related or unrelated to those elementsspecifically identified unless clearly indicated to the contrary. Thus,as a non-limiting example, a reference to “A and/or B,” when used inconjunction with open-ended language such as “comprising” can refer, inone embodiment, to A without B (optionally including elements other thanB); in another embodiment, to B without A (optionally including elementsother than A); in yet another embodiment, to both A and B (optionallyincluding other elements); etc.

The word “or” should be understood to have the same meaning as “and/or”as defined above. For example, when separating items in a list, “or” or“and/or” shall be interpreted as being inclusive, i.e., the inclusion ofat least one, but also including more than one, of a number or list ofelements, and, optionally, additional unlisted items. Only terms clearlyindicated to the contrary, such as “only one of” or “exactly one of,” ormay refer to the inclusion of exactly one element of a number or list ofelements. In general, the term “or” as used herein shall only beinterpreted as indicating exclusive alternatives (i.e. “one or the otherbut not both”) when preceded by terms of exclusivity, such as “either,”“one of,” “only one of,” or “exactly one of.”

The phrase “at least one,” in reference to a list of one or moreelements, should be understood to mean at least one element selectedfrom any one or more of the elements in the list of elements, but notnecessarily including at least one of each and every elementspecifically listed within the list of elements and not excluding anycombinations of elements in the list of elements. This definition alsoallows that elements may optionally be present other than the elementsspecifically identified within the list of elements to which the phrase“at least one” refers, whether related or unrelated to those elementsspecifically identified. Thus, as a non-limiting example, “at least oneof A and B” (or, equivalently, “at least one of A or B,” or,equivalently “at least one of A and/or B”) can refer, in one embodiment,to at least one, optionally including more than one, A, with no Bpresent (and optionally including elements other than B); in anotherembodiment, to at least one, optionally including more than one, B, withno A present (and optionally including elements other than A); in yetanother embodiment, to at least one, optionally including more than one,A, and at least one, optionally including more than one, B (andoptionally including other elements); etc. The transitional words orphrases, such as “comprising,” “including,” “carrying,” “having,”“containing,” “involving,” “holding,” and the like, are to be understoodto be open-ended, i.e., to mean including but not limited to.

The terms “pyrolysis” and “pyrolyzing” refer to the transformation of amaterial (e.g., a solid hydrocarbonaceous material) into one or moreother materials (e.g., volatile organic compounds, gases, coke, etc.) byheat, without oxygen or other oxidants or without significant amounts ofoxygen or other oxidants, and with or without the use of a catalyst.

The term “catalytic pyrolysis” refers to pyrolysis performed in thepresence of a catalyst.

The terms “aromatics” or “aromatic compound” refer to a hydrocarboncompound or compounds comprising one or more aromatic groups such as,for example, single aromatic ring systems (e.g., benzyl, phenyl, etc.)and/or fused polycyclic aromatic ring systems (e.g. naphthyl,1,2,3,4-tetrahydronaphthyl, etc.). Examples of aromatic compoundsinclude, but are not limited to, benzene, toluene, indane, indene,2-ethyl toluene, 3-ethyl toluene, 4-ethyl toluene, trimethyl benzene(e.g., 1,3,5-trimethyl benzene, 1,2,4-trimethyl benzene, 1,2,3-trimethylbenzene, etc.), ethylbenzene, styrene, cumene, methylbenzene,propylbenzene, xylenes (e.g., p-xylene, m-xylene, o-xylene, etc.),naphthalene, methyl-naphthalene (e.g., 1-methyl naphthalene, anthracene,9,10-dimethylanthracene, pyrene, phenanthrene, dimethyl-naphthalene(e.g., 1,5-dimethylnaphthalene, 1,6-dimethylnaphthalene,2,5-dimethylnaphthalene, etc.), ethyl-naphthalene, hydrindene,methyl-hydrindene, and dymethyl-hydrindene. Single ring and/or higherring aromatics may be produced in some embodiments.

The term “petrochemicals” is used herein to refer to chemicals, chemicalprecursors, chemical intermediates, and the like, traditionally derivedfrom petroleum sources. Petrochemicals include paraffins, olefins,aromatic compounds, and the like. For purposes of this application, whenthese materials are derived from biomass, as well as other non-petroleumsources (e.g., recycled plastics, municipal solid waste, sugar canebagasse, wood, etc.), the term petrochemicals may be employed despitethe fact that the chemicals, chemical precursors, chemicalintermediates, and the like, may not be derived directly from petroleum.

The term “biomass” refers to living and recently dead biologicalmaterial. In accordance with the inventive method, biomass may beconverted, for example, to liquid fuel (e.g., biofuel or biodiesel) orto other fluid hydrocarbon products. Biomass may include trees (e.g.,wood) as well as other vegetation; agricultural products andagricultural waste (e.g., corn stover, bagasse, fruit, garbage, silage,etc.); energy crops (e.g. switchgrass, miscanthus); algae and othermarine plants; metabolic wastes (e.g., manure, sewage); and cellulosicurban waste. Biomass may be considered as comprising material thatrecently participated in the carbon cycle so that the release of carbonin a combustion process may result in no net increase averaged over areasonably short period of time. For this reason, peat, lignite, coal,shale oil or petroleum may not be considered as being biomass as theycontain carbon that may not have participated in the carbon cycle for along time and, as such, their combustion may result in a net increase inatmospheric carbon dioxide. The term biomass may refer to plant mattergrown for use as biofuel, but may also includes plant or animal matterused for production of fibers, chemicals, heat, and the like. Biomassmay also include biodegradable waste or byproducts that can be burnt asfuel or converted to chemicals. These may include municipal waste, greenwaste (the biodegradable waste comprised of garden or park waste such asgrass or flower cuttings, hedge trimmings, and the like), byproducts offarming including animal manures, food processing wastes, sewage sludge,black liquor from wood pulp or algae, and the like. Biomass may bederived from plants, including miscanthus, spurge, sunflower,switchgrass, hemp, corn (maize), poplar, willow, sugarcane, and oil palm(palm oil), and the like. Biomass may be derived from roots, stems,leaves, seed husks, fruits, and the like. The particular plant or otherbiomass source used may not be important to the fluid hydrocarbonproduct produced in accordance with the inventive method, although theprocessing of the biomass may vary according to the needs of the reactorand the form of the biomass.

The hydrocarbonaceous feed material for the inventive method maycomprise a solid hydrocarbonaceous material, a semi-solidhydrocarbonaceous material, a liquid hydrocarbonaceous material, or amixture of two or more thereof. The solids content of thehydrocarbonaceous feed may be up to about 100% by weight, or from about30% to about 100% by weight, or from about 50% to about 100%, or fromabout 70% to about 100%, or from 90% to about 100%, or from about 95% toabout 100%, or from about 98% to about 100%, or from about 30% to about95%, or from about 50% to about 95%, or from about 70% to about 95%, orfrom about 80% to about 95%, or from about 85% to about 95%, or fromabout 90% to about 95% by weight. The hydrocarbonaceous material maycomprise biomass. The hydrocarbonaceous material may comprise plasticwaste, recycled plastics, agricultural solid waste, municipal solidwaste, food waste, animal waste, carbohydrates, lignocellulosicmaterials, xylitol, glucose, cellobiose, hemi-cellulose, lignin, sugarcane bagasse, glucose, wood, corn stover, or a mixture of two or morethereof. The hydrocarbonaceous material may comprise furan and/or2-methylfuran. The hydrocarbonaceous material may comprise pinewood. Thehydrocarbonaceous material may comprise pyrolysis oil derived frombiomass, a carbohydrate derived from biomass, an alcohol derived frombiomass, a biomass extract, a pretreated biomass, a digested biomassproduct, or a mixture of two or more thereof. Mixtures of two or more ofany of the foregoing may be used.

The inventive method may comprise feeding the hydrocarbonaceous materialto a reactor. At least a portion of the hydrocarbonaceous material maybe pyrolyzed in the reactor under reaction conditions sufficient toproduce one or more pyrolysis products. At least a portion of thepyrolysis products may be catalytically reacted under sufficientconditions to produce the fluid hydrocarbon product. The reactor maycomprise a continuously stirred tank reactor, a batch reactor, asemi-batch reactor, a fixed bed reactor, or a fluidized bed reactor.Advantageously, the reactor may comprise a fluidized bed reactor. Thecatalytic reaction step may be achieved by co-feeding the catalyst withthe hydrocarbonaceous material. The catalyst may be fed separately. Partof the catalyst may be fed with the hydrocarbonaceous feed material andpart of the catalyst may be fed separately.

The inventive method may be used for the production of fluid (e.g., aliquid, a supercritical fluid, and/or a gas) hydrocarbon products via acatalytic pyrolysis process (e.g., a CFP process). The fluid hydrocarbonproduct, or a portion thereof, may comprise a liquid at standard ambienttemperature and pressure (SATP—i.e. 25° C. and 100 kPa absolutepressure). The hycrocarbonaceous feed material may be pyrolyzed atintermediate temperatures (for example, in the range from about 400° C.and about 600° C.), compared to temperatures typically used in the priorart. The pyrolysis step may be conducted for an effective amount of timeto produce discrete, identifiable fluid hydrocarbon products. Theinventive method may involve heating the hydrocarbonaceous material (andoptionally the catalyst) to the reaction temperature at relatively highheating rates (e.g., greater than about 50° C. per second).

The inventive method may involve the use of specialized catalysts. Thesecatalysts are zeolite catalysts which contain silica and alumina. Thecatalyst may be characterized by pores with pore mouth openings whereinthe pore mouth openings have been reduced in size, and catalytic siteson the external surface of the catalyst that have been covered orobscured. The catalyst may have catalytic sites in the pores near thepore mouth openings that have been covered or obscured. The catalyst maybe treated with a silicone compound to reduce the size of the pore mouthopenings and cover or obscure catalytic sites on the external surface ofthe catalyst and in the pores of the catalyst near the pore mouthopenings. The catalyst may comprise relatively small particles, whichmay be agglomerated. The composition fed to the reactor may have arelatively high catalyst to hydrocarbonaceous material mass ratio (e.g.,from about 0.33:1 to about 20:1, or from about 0.5:1 to about 20:1, orfrom about 2:1 to about 10:1).

The expression that catalytic sites are positioned “in the pores nearthe pore mouth openings” refers to catalytic sites within the pores thatare rendered inaccessible to the pyrolysis product as a result oftreatment with the silicone compound. In an embodiment, these catalyticsites may be positioned within the pores at a depth of no more thanabout 10 angstroms from the pore mouth openings, or no more than about 7angstroms, or no more than about 5 angstroms, or no more than about 2angstroms, from the pore mouth openings. These catalytic sites may beinhibited or inoperative as a result of treatment with the siliconecompound. That is, these catalysts may be deactivated as a result of thetreatment with the silicone compound.

The inventive method may comprise a single-stage method for thepyrolysis of the hydrocarbonaceous material. This method may compriseproviding or using a single-stage pyrolysis apparatus. A single-stagepyrolysis apparatus may be one in which pyrolysis and subsequentcatalytic reactions are carried out in a single vessel. The single-stagepyrolysis apparatus may comprise a continuously stirred tank reactor, abath reactor, a semi-batch reactor, a fixed bed reactor or a fluidizedbed reactor. Multi-stage apparatuses may also be used for the productionof fluid hydrocarbon products in accordance with the invention.

The hydrocarbonaceous material may comprise solids of any suitable size.In some cases, it may be advantageous to use hydrocarbonaceous solidswith relatively small particle sizes. Small-particle solids may, in someinstances, react more quickly than larger solids due to their relativelyhigher surface area to volume ratios compared to larger solids. Inaddition, small particle sizes may allow for more efficient heattransfer within each particle and/or within the reactor volume. This mayprevent or reduce the formation of undesired reaction products.Moreover, small particle sizes may provide for increased solid-gas andsolid-solid contact, leading to improved heat and mass transfer. Theaverage particle size of the solid hydrocarbonaceous material may beless than about 5 mm, less than about 2 mm, less than about 1 mm, lessthan about 500 microns, less than about 60 mesh (250 microns), less thanabout 100 mesh (149 microns), less than about 140 mesh (105 microns),less than about 170 mesh (88 microns), less than about 200 mesh (74microns), less than about 270 mesh (53 microns), or less than about 400mesh (37 microns), or smaller.

It may be desirable to employ a feed material with an average particlesize above a minimum amount in order to reduce the pressure required topass the solid hydrocarbonaceous feed material through the reactor. Forexample, it may be desirable to use a solid hydrocarbonaceous feedmaterial with an average particle size of at least about 400 mesh (37microns), at least about 270 mesh (53 microns), at least about 200 mesh(74 microns), at least about 170 mesh (88 microns), at least about 140mesh (105 microns), at least about 100 mesh (149 microns), at leastabout 60 mesh (250 microns), at least about 500 microns, a least about 1mm, at least about 2 mm, at least about 5 mm, or higher.

The hydrocarbonaceous material may comprise biomass. Thehydrocarbonaceous material may comprise plastic waste, recycledplastics, agricultural and/or municipal solid waste, food waste, animalwaste, carbohydrates, lignocellulosic materials (e.g., wood chips orshavings), or a mixture of two or more thereof. The hydrocarbonaceousmaterial may comprise xylitol, glucose, cellobiose, cellulose,hemi-cellulose, lignin, or a mixture of two or more thereof. Thehydrocarbonaceous material may comprise sugar cane bagasse, glucose,wood, corn stover, or a mixture of two or more thereof. Thehydrocarbonaceous material may comprise wood.

Biomass pyrolysis liquid or bio-oil may be formed during the pyrolyzingstep of the inventive method. Biomass pyrolysis liquid may be dark brownand may approximate to biomass in elemental composition. It may becomposed of a very complex mixture of oxygenated hydrocarbons with anappreciable proportion of water from both the original moisture andreaction product. Compositionally, biomass pyrolysis oil may vary withthe type of biomass, but is known to contain oxygenated low molecularweight alcohols (e.g., furfuryl alcohol), aldehydes (aromaticaldehydes), ketones (furanone), phenols (methoxy phenols) and water.Solid char may also be present, suspended in the oil. The liquid may beformed by rapidly quenching the intermediate products of flash pyrolysisof hemicellulose, cellulose and lignin in the biomass. Chemically, theoil may contain several hundred different chemicals in widely varyingproportions, ranging from formaldehyde and acetic acid to complex highmolecular weight phenols, anhydrosugars and other oligosaccharides. Itmay have a distinctive odor from low molecular weight aldehydes andacids, and may be acidic with a pH of about 1.5 to about 3.8, and can bean irritant.

The residence time of the catalyst in the reactor may be defined as thevolume of the reactor filled with catalyst divided by the volumetricflow rate of the catalyst through the reactor. For example if a 3 literreactor contains 2 liters of catalyst and a flow of 0.4 liters perminute of catalyst is fed through the reactor, i.e., both fed andremoved, the catalyst residence time will be 2/0.4 minutes, or 5minutes.

The residence time of the catalyst in the reactor may be at least about1 minute, at least about 2 minutes, at least about 5 minutes, at leastabout 7 minutes, at least about 10 minutes, at least about 15 minutes,at least about 20 minutes, at least about 25 minutes, at least about 30minutes, at least about 60 minutes, or at least about 120 minutes. Insome cases, the residence time of the catalyst in the reactor may beless than about 120 minutes, or from about 1 minute and about 120minutes, or from about 2 minutes to about 120 minutes, or from about 5minutes to about 120 minutes, or from about 7 minutes to about 120minutes, or from about 10 minutes to about 120 minutes, or from about 12minutes to about 120 minutes, or from about 15 minutes to about 120minutes, or from about 20 minutes to about 120 minutes, or from about 30minutes to about 120 minutes, or from about 60 minutes to about 120minutes. In some cases, the use of relatively long residence times mayallow for additional chemical reactions to form desirable products. Longcatalyst residence times may be achieved by, for example, increasing thevolume of the reactor and/or reducing the volumetric flow rate of thecatalyst. The residence time of the catalyst may be relatively short,e.g., less than about 120 minutes, or less than about 60 minutes.

Contact time may be defined as the residence time of a material in areactor or other device, when measured or calculated under standardconditions of temperature and pressure (i.e., 0° C. and 100 kPa absolutepressure). For example, a 2 liter reactor to which is fed 3 standardliters per minute of gas has a contact time of ⅔ minute, or 40 secondsfor that gas. For a chemical reaction, contact time or residence time isbased on the volume of the reactor where substantial reaction isoccurring; and would exclude volume where substantially no reaction isoccurring such as an inlet or an exhaust conduit. For catalyzedreactions, the volume of a reaction chamber is the volume where catalystis present.

The term “conversion of a reactant” may refer to the reactant mole ormass change between a material flowing into a reactor and a materialflowing out of the reactor divided by the moles or mass of reactant inthe material flowing into the reactor. For example, if 100 g of ethyleneare fed to a reactor and 30 g of ethylene are flowing out of thereactor, the conversion is [(100−30)/100]=70% conversion of ethylene.

The term “fluid” may refer to a gas, a liquid, a mixture of a gas and aliquid, or a gas or a liquid containing dispersed solids, liquiddroplets and/or gaseous bubbles. The terms “gas” and “vapor” have thesame meaning and are sometimes used interchangeably. In someembodiments, it may be advantageous to control the residence time of thefluidization fluid in the reactor. The fluidization residence time ofthe fluidization fluid is defined as the volume of the reactor dividedby the volumetric flow rate of the fluidization fluid under processconditions of temperature and pressure.

The term “fluidized bed reactor” may be used to refer to reactorscomprising a vessel that contains a granular solid material (e.g.,silica particles, catalyst particles, etc.), in which a fluid (e.g., agas or a liquid) is passed through the granular solid material atvelocities sufficiently high as to suspend the solid material and causeit to behave as though it were a fluid. The term “circulating fluidizedbed reactor” may be used to refer to fluidized bed reactors in which thegranular solid material is passed out of the reactor, circulated througha line in fluid communication with the reactor, and recycled back intothe reactor.

Bubbling fluidized bed reactors and turbulent fluidized bed reactors maybe used. In bubbling fluidized bed reactors, the fluid stream used tofluidize the granular solid material may be operated at a sufficientlylow flow rate such that bubbles and voids may be observed within thevolume of the fluidized bed during operation. In turbulent fluidized bedreactors, the flow rate of the fluidizing stream will be higher thanthat employed in a bubbling fluidized bed reactor. Examples of fluidizedbed reactors, circulating fluidized bed reactors, bubbling and turbulentfluidized bed reactors are described in Kirk-Othmer Encyclopedia ofChemical Technology (online), Vol. 11, Hoboken, N.J.: WileyInterscience, 2001, pages 791-825, these pages being incorporated hereinby reference.

The terms “olefin” or “olefin compound” (a.k.a. “alkenes”) may be usedto refer to any unsaturated hydrocarbon containing one or more pairs ofcarbon atoms linked by a double bond. Olefins may include both cyclicand acyclic (aliphatic) olefins, in which the double bond is locatedbetween carbon atoms forming part of a cyclic (closed-ring) or of anopen-chain grouping, respectively. In addition, olefins may include anysuitable number of double bonds (e.g., monoolefins, diolefins,triolefins, etc.). Examples of olefin compounds may include ethene,propene, allene (propadiene), 1-butene, 2-butene, isobutene (2 methylpropene), butadiene, and isoprene, among others. Examples of cyclicolefins may include cyclopentene, cyclohexane, cycloheptene, amongothers. Aromatic compounds such as toluene are not considered olefins;however, olefins that include aromatic moieties are considered olefins,for example, benzyl acrylate or styrene.

Pore size relates to the size of a molecule or atom that can penetrateinto the pores of a material. As used herein, the term “pore size” forzeolites refers to the Norman radii adjusted pore size. Determination ofNorman radii adjusted pore size is described, for example, in Cook, M.;Conner, W. C., “How big are the pores of zeolites?” Proceedings of theInternational Zeolite Conference, 12th, Baltimore, Jul. 5-10, 1998;(1999), 1, pp 409-414, which is incorporated herein by reference. As aspecific exemplary calculation, the atomic radii for ZSM-5 pores areabout 5.5-5.6 Angstroms, as measured by x-ray diffraction. In order toadjust for the repulsive effects between the oxygen atoms in thecatalyst, Cook and Conner have shown that the Norman adjusted radii are0.7 Angstroms larger than the atomic radii (about 6.2-6.3 Angstroms).

One of ordinary skill in the art will understand how to determine thepore size (e.g., minimum pore size, average of minimum pore sizes) in acatalyst. For example, x-ray diffraction (XRD) may be used to determineatomic coordinates. XRD techniques for the determination of pore sizeare described, for example, in Pecharsky, V. K. et al, “Fundamentals ofPowder Diffraction and Structural Characterization of Materials,”Springer Science+Business Media, Inc., New York, 2005, incorporatedherein by reference in its entirety. Other techniques that may be usefulin determining pore sizes (e.g., zeolite pore sizes) may include, forexample, helium pycnometry or low pressure argon adsorption techniques.These and other techniques are described in Magee, J. S. et at, “FluidCatalytic Cracking: Science and Technology,” Elsevier PublishingCompany, Jul. 1, 1993, pp. 185-195, which is incorporated herein byreference in its entirety. Pore sizes of mesoporous catalysts may bedetermined using, for example, nitrogen adsorption techniques, asdescribed in Gregg, S. J. at al, “Adsorption, Surface Area andPorosity,” 2nd Ed., Academic Press Inc., New York, 1982 and Rouquerol,F. et al, “Adsorption by powders and porous materials. Principles,Methodology and Applications,” Academic Press Inc., New York, 1998, bothof which are incorporated herein by reference.

A screening method may be used to select catalysts with appropriate poresizes for the conversion of specific pyrolysis product molecules. Thescreening method may comprise determining the size of pyrolysis productmolecules desired to be catalytically reacted (e.g., the moleculekinetic diameters of the pyrolysis product molecules). One of ordinaryskill in the art may calculate, for example, the kinetic diameter of agiven molecule. The type of catalyst may then be chosen such that thepores of the catalyst (e.g., Norman adjusted minimum radii) aresufficiently large to allow the pyrolysis product molecules to diffuseinto and/or react with the catalyst. In some embodiments, the catalystsmay be chosen such that their pore sizes are sufficiently small toprevent entry and/or reaction of pyrolysis products whose reaction wouldbe undesirable.

The catalyst may comprise any catalyst suitable for conducting thecatalytically reacting step of the inventive method. The catalyst may beused to lower the activation energy (increase the rate) of the reactionconducted in the catalytically reacting step and/or improve thedistribution of products or intermediates during the reaction (forexample, a shape selective catalyst). Examples of reactions that can becatalyzed include: dehydration, dehydrogenation, isomerization, hydrogentransfer, aromatization, decarbonylation, decarboxylation, aldolcondensation, and combinations thereof. The catalyst components may beacidic, neutral or basic.

The inventive method may comprise a CFP process. For CFP processes,particularly advantageous catalysts may include those containinginternal porosity selected according to pore size (e.g., mesoporous andpore sizes typically associated with zeolites), e.g., average pore sizesof less than about 100 Angstroms, less than about 50 Angstroms, lessthan about 20 Angstroms, less than about 10 Angstroms, less than about 5Angstroms, or smaller. In some embodiments, catalysts with average poresizes of from about 5 Angstroms to about 100 Angstroms may be used. Insome embodiments, catalysts with average pore sizes of between about 5.5Angstroms and about 6.5 Angstroms, or between about 5.9 Angstroms andabout 6.3 Angstroms may be used. In some cases, catalysts with averagepore sizes of between about 7 Angstroms and about 8 Angstroms, orbetween about 7.2 Angstroms and about 7.8 Angstroms may be used.

The catalyst may be selected from naturally occurring zeolites,synthetic zeolites and combinations thereof. The catalyst may comprise aZSM-5 zeolite catalyst. The catalyst may comprise acid sites. These acidsites may also be referred to as catalytically active sites. Otherzeolite catalysts that may be used may include ferrierite, zeolite Y,zeolite beta, mordenite, MCM-22, ZSM-23, ZSM-57, SUZ-4, EU-1, ZSM-11,(S)AlP0-31, SSZ-23, and the like. The catalyst may comprise silica andalumina, and further comprise one or more additional metals and/or ametal oxides. Suitable metals and/or oxides may include, for example,nickel, palladium, platinum, titanium, vanadium, chromium, manganese,iron, cobalt, zinc, copper, gallium, and/or any of their oxides, amongothers. In some cases promoter elements selected from the rare earthelements, i.e., elements 57-71, cerium, zirconium or their oxides, orcombinations of these may be included to modify the activity, structureand/or stability of the catalyst. In addition, in some cases, propertiesof the catalysts (e.g., pore structure, type and/or number of catalyticsites, etc.) may be chosen to selectively produce a desired product.

The catalyst may be treated or impregnated one or more times with asilicone compound to reduce the size of the pore mouth openings in thecatalyst as well as cover or obscure catalytic sites on the externalsurface of the catalyst and inside the pores of the catalyst near thepore mouth openings. The covering of the catalytic sites with thetreatment layer may inhibit and/or extinguish their catalytic activity.In order to facilitate a more controlled application of the siliconecompound, the silicone compound may be dispersed in a carrier, forexample, an aqueous or organic liquid carrier.

In each phase of the catalyst treatment process, the silicone compoundmay be deposited on the external surface of the catalyst by any suitablemethod. For example, the silicone compound may be dissolved in anorganic carrier, mixed with the catalyst, and then dried by evaporationor vacuum distillation. The catalyst may be contacted with the siliconecompound at a catalyst to silicone compound weight ratio in the rangefrom about 1000:1 to about 1:10.

The silicone compound may be provided in the form of a solution or anemulsion under the conditions of contact with the catalyst. Thedeposited silicone compound may cover, and reside substantiallyexclusively on, the external surface of the catalyst, blocking externalsites and partially blocking pore mouths and sites in or near the poremouths openings. Examples of methods of depositing silicone compounds onthe surface of zeolites may be found in U.S. Pat. Nos. 4,090,981;5,243,117; 5,403,800, and 5,659,098, which are incorporated by referenceherein.

The catalyst may be ex situ treated by single or multiple coatings withthe silicone compound, each coating followed by calcination. The coatedcatalyst may comprise silica and alumina. The silica to alumina molarratio may be in the range from about 10:1 to about 50:1, or in the rangefrom about 10:1 to about 40:1, or in the range from about 10:1 to about20:1, or about 15:1. The coated catalyst may further comprise nickel,platinum, vanadium, palladium, manganese, cobalt, zinc, copper,chromium, gallium, an oxide of one or more thereof, or a mixture of twoor more thereof.

The term “silicone compound” is used herein to refer to any compoundthat contains one or more Si—O groups. The silicone compound may be asilicate containing one or more of SiO₄ ⁴⁻, Si₂O₇ ⁶⁻ or Si₆O₁₈ ¹²⁻groups. These may include one or more tetraorthosilicates. The siliconecompound may include one or more siloxanes containing one or moresilicon-oxygen backbones (—Si—O—Si—O—) with organic (e.g., hydrocarbon)side groups attached to the silicon atoms. These may include one or moresiloxane polymers (e.g., polydimethyl siloxane). The silicone compoundmay be a straight chain, branched chain or cyclical compound. Thesilicone compound may be monomeric, oligomeric or polymeric. Thesilicone compound may comprise a compound containing at least one grouprepresented by the formula

The silicone compound may be represented by the formula:

wherein R₁ and R₂ independently comprise hydrogen, halogen, hydroxyl,alkyl, alkoxyl, halogenated alkyl, aryl, halogenated aryl, aralkyl,halogenated aralkyl, alkaryl or halogenated alkaryl; and n is a numberthat is at least 2. R₁ and/or R₂ may comprise methyl, ethyl or phenyl. nmay be a number in the range from about 3 to about 1000.

The silicone compound may have a number average molecular weight in therange from about 80 to about 20,000, or from about 150 to 10,000.

The silicone compound may comprise dimethylsilicone, diethylsilicone,phenylmethylsilicone, methylhydrogensilicone, ethylhydrogen silicone,phenylhydrogen silicone, methylethyl silicone, phenylethyl silicone,diphenyl silicone, methyltrifluoropropyl silicone, ethyltrifluoropropylsilicone, polydimethyl silicone, tetrachloro-phenylmethyl silicone,tetrachlorophenylethyl silicone, tetrachlorophenylhydrogen silicone,tetrachlorophenylphenyl silicone, methylvinyl silicone, hexamethylcyclotrisiloxane, octamethyl cyclotetrasiloxane, hexaphenylcyclotrisiloxane, octaphenyl cyclotetrasiloxane, or a mixture of two ormore thereof.

The silicone compound may comprise a tetraorthosilicate. The siliconecompound may comprise tetramethylorthosilicate, tetraethylorthosilicate,or a mixture thereof.

The kinetic diameter of the silicone compound may be larger than thepore diameter of the catalyst in order to avoid entry of the siliconecompound into the pore and any concomitant reduction in the internalactivity of the catalyst.

The organic carrier for the silicone compound may comprise hydrocarbonssuch as linear, branched, and cyclic alkanes having five or morecarbons. The carrier may comprise a linear, branched or cyclic alkanehaving a boiling point greater than about 70° C., and containing about 6or more carbons. Optionally, mixtures of low volatility organiccompounds, such as hydrocracker recycle oil, may also be employed ascarriers. Low volatility hydrocarbon carriers for the silicone compoundmay comprise decane, dodecane, mixtures thereof, and the like.

Following each deposition of the silicone compound, the catalyst may becalcined to decompose the molecular or polymeric species to a solidstate species. The catalyst may be calcined at a rate of from about 0.2°C./minute to about 5° C./minute to a temperature greater than about 200°C., but below a temperature at which the crystallinity of the zeolitemay be adversely affected. Generally, such temperature will be belowabout 600° C. The temperature of calcination may be in the range fromabout 350° C. to about 550° C. The catalyst may be maintained at thecalcination temperature for about 1 to about 24 hours, or about 2 toabout 6 hours.

The catalyst may be treated with a tetraorthosilicate using a chemicalliquid deposition (CLD) process. The tetraorthosilicate may comprisetetramethylorthosilicate, tetraethylorthosilicate (TEOS), or a mixturethereof. The CLD process may comprise dispersing the catalyst in aliquid medium (e.g., hexanes, alkanes, aromatics or other non-polarorganic solvent) at a concentration in the range from about 0.1 to about20% by weight, or from about 1 to about 10% by weight; adding thetetraorthosilicate to provide a mixture containing from about 0.01 toabout 5% by weight, or from about 0.1 to about 1% by weight of thetetraorthosilicate; refluxing the resulting mixture at an elevatedtemperature (e.g., in the range from about 50 to about 150° C., or about90° C.) with stirring; recovering the catalyst (e.g., via centrifuging)from the liquid medium; drying the catalyst; and then calcining thecatalyst in air at a temperature in the range from about 100 to about550° C., or from about 150 to about 325° C., for a time period in therange from about 1 to about 24 hours, or about 2 to about 12 hours. Thisprocedure may be repeated any desired number of times (e.g., 1, 2, 3additional times, etc.), to provide for the desired treatment layerderived from the tetraorthosilicate. When using TEOS as thetetraorthosilicate, this process may be referred to as a TEOS CLDsilylation process.

While not wishing to be bound by theory, it is believed that theadvantages of treatment with silicone compounds are in part obtained byrendering active catalytic sites on the external surfaces of thecatalyst substantially inaccessible to reactants, while increasingcatalyst tortuosity by reducing the size of the pore mouth openings.Active catalytic sites existing on the external surface of the catalystare believed to isomerize the para-isomer back to an equilibrium levelwith the other two isomers. Thus, by reducing the availability of theseactive catalytic sites, the relatively high proportion of para-xylenemay be maintained. It is believed that the silicone compounds of thepresent invention may block or otherwise render these external catalyticsites unavailable to the para-isomers by chemically modifying, covering,or obscuring the sites.

It may be beneficial to control the residence time of the reactants(e.g., the solid hydrocarbonaceous material and/or a non-solid reactant)and catalyst(s) in the reactor and/or under a defined set of reactionconditions (i.e. conditions under which the reactants may undergopyrolysis or catalysis in a given reactor system).

The term “overall residence time” refers to the volume of a reactor ordevice or specific portion of a reactor or device divided by the exitflow of all gases out of the reactor or device including fluidizationgas, products, and impurities, measured or calculated at the averagetemperature of the reactor or device and the exit pressure of thereactor or device.

The term “reactant residence time” of a reactant in the reactor isdefined as the amount of time the reactant spends in the reactor.Residence time may be based on the feed rate of reactant and isindependent of rate of reaction. The reactant residence time of thereactants in a reactor may be calculated using different methodsdepending upon the type of reactor being used. For gaseous reactants,where flow rate into the reactor is known, this is typically a simplecalculation. In the case of solid reactants in which the reactorcomprises a packed bed reactor into which only reactants arecontinuously fed (i.e. no carrier or fluidizing flow is utilized), thereactant residence time in the reactor may be calculated by dividing thevolume of the reactor by the volumetric flow rate of thehydrocarbonaceous material and fluid hydrocarbon product exiting thereactor.

In cases where the reaction takes place in a reactor that is closed tothe flow of mass during operation (e.g., a batch reactor), the batchresidence time of the reactants in such may be reactor is defined as theamount of time elapsing between the time at which the temperature in thereactor containing the reactants reaches a level sufficient to commencea pyrolysis reaction (e.g., for CFP, typically about 300° C. to about1000° C. for many typical hydrocarbonaceous feedstock materials) and thetime at which the reactor is quenched (e.g., cooled to a temperaturebelow that sufficient to support further pyrolysis—e.g. typically about300° C. to about 1000° C. for many hydrocarbonaceous feedstockmaterials).

In some cases, e.g. for certain fluidized bed reactors, the reactor feedstream(s) may include feed stream(s) comprising auxiliary materials(i.e., matter other than solid hydrocarbonaceous materials and/ornon-solid reactants). For example, in certain cases where fluidized bedsare used as reactors, the feed stream may comprise fluidizationfluid(s). In cases where circulating fluidized beds are used, catalystand fluidization fluid may both be fed, recycled, or fed and recycled tothe reactor. In such cases, the reactant residence time of the reactantsin the reactor can be determined as the volume of the reactor divided bythe volumetric flow rate of the reactants and reaction product gasesexiting the reactor as with the packed bed situation described above;however, since the flow rate of the reactants and reaction product gasesexiting the reactor may not be convenient to determine directly, thevolumetric flow rate of the reactants and reaction product gases exitingthe reactor may be estimated by subtracting the feed volumetric flowrate of the auxiliary materials (e.g., fluidization fluid, catalyst,contaminants, etc.) into the reactor from the total volumetric flow rateof the gas stream(s) exiting the reactor.

The term “selectivity” refers to the amount of production of aparticular product in comparison to a selection of products. Selectivityto a product may be calculated by dividing the amount of a particularproduct by the amount of a number of products produced. For example, if75 grams of aromatics are produced in a reaction and 20 grams of benzeneare found in these aromatics, on a mass basis the selectivity to benzeneamongst aromatic products is 20/75=26.7%. Selectivity may be calculatedon a mass basis, as in the aforementioned example, or it may becalculated on a carbon basis where the selectivity is calculated bydividing the amount of carbon that is found in a particular product bythe amount of carbon that is found in a selection of products. Unlessspecified otherwise, for reactions involving biomass as a reactant,selectivity is on a mass basis. For reactions involving conversion of aspecific molecular reactant (ethene for example), selectivity is thepercentage (on a mass basis unless specified otherwise) of a selectedproduct divided by all the products produced. The selectivity forvarious materials can be determined using the following equations:

$\begin{matrix}{{{Overall}\mspace{14mu} {selectivity}} = {\frac{{moles}\mspace{14mu} {of}\mspace{14mu} {carbon}\mspace{14mu} {in}\mspace{14mu} a\mspace{14mu} {product}}{{moles}\mspace{14mu} {of}\mspace{14mu} {carbon}\mspace{14mu} {in}\mspace{14mu} {all}\mspace{14mu} {products}} \times 100\%}} & (1) \\{{{Aromatic}\mspace{14mu} {selectivity}} = {\frac{{moles}\mspace{14mu} {of}\mspace{14mu} {carbon}\mspace{14mu} {in}\mspace{14mu} {an}\mspace{14mu} {aromatic}\mspace{14mu} {product}}{{moles}\mspace{14mu} {of}\mspace{14mu} {carbon}\mspace{14mu} {in}\mspace{14mu} {all}\mspace{14mu} {aromatic}\mspace{14mu} {products}} \times 100\%}} & (2) \\{{{Olefin}\mspace{14mu} {selectivity}} = {\frac{{moles}\mspace{14mu} {of}\mspace{14mu} {carbon}\mspace{14mu} {in}\mspace{14mu} {an}\mspace{14mu} {olefinic}\mspace{14mu} {product}}{{moles}\mspace{14mu} {of}\mspace{14mu} {carbon}\mspace{14mu} {in}\mspace{14mu} {all}\mspace{14mu} {olefins}\mspace{14mu} {products}} \times 100\%}} & (3) \\{{p\text{-}{Xylene}\mspace{14mu} {selectivity}\mspace{14mu} {in}\mspace{14mu} {xylenes}} = {\frac{{moles}\mspace{14mu} {of}{\mspace{11mu} \;}p\text{-}{xylene}\mspace{14mu} {isomer}}{{moles}\mspace{14mu} {of}\mspace{14mu} {all}\mspace{14mu} {xylene}\mspace{14mu} {isomers}} \times 100\%}} & (4)\end{matrix}$

The term “yield” is used herein to refer to the amount of a productflowing out of a reactor divided by the amount of reactant flowing intothe reactor, usually expressed as a percentage or fraction. Yields areoften calculated on a mass basis, carbon basis, or on the basis of aparticular feed component. Mass yield is the mass of a particularproduct divided by the weight of feed used to prepare that product. Forexample, if 500 grams of biomass is fed to a reactor and 45 grams ofp-xylene is produced, the mass yield of p-xylene would be 45/500=9%p-xylene. Carbon yield is the mass of carbon found in a particularproduct divided by the mass of carbon in the feed to the reactor. Forexample, if 500 grams of biomass that contains 40% carbon is reacted toproduce 45 g of p-xylene that contains 90.6% carbon, the carbon yield is[(45*0.906)/(500*0.40)]=20.4%. Carbon yield from biomass is the mass ofcarbon found in a particular product divided by the mass of carbon fedto the reactor in a particular feed component. For example, if 500 gramsof biomass containing 40% carbon and 100 grams of CO₂ are reacted toproduce 40 g of p-xylene (containing 90.6% carbon), the carbon yield onbiomass is [(40*0.906)/(500*0.40)]=18.1%; note that the mass of CO₂ doesnot enter into the calculation.

The embodiments described herein may also involve chemical processdesigns used to perform catalytic pyrolysis. The processes may involvethe use of one or more fluidized bed reactors (e.g., a circulatingfluidized bed reactor, turbulent fluidized bed reactor, bubblingfluidized bed reactor, etc.). The process designs described herein mayoptionally involve specialized handling of the material fed to one ormore reactors. For example, the feed material may be dried, cooled,and/or ground prior to supplying the material to a reactor. Otheraspects of the invention may relate to product compositions producedusing the process designs described herein.

Without being bound to a particular mode of action or order of steps ofthe overall thermal/catalytic conversion process, catalytic pyrolysis isbelieved to involve at least partial thermal pyrolysis ofhydrocarbonaceous material (e.g., solid biomass such as cellulose) toproduce one or more pyrolysis products (e.g., volatile organics, gases,solid coke, etc.) and catalytic reaction of at least a portion of theone or more pyrolysis products using a catalyst under reactionconditions sufficient to produce fluid hydrocarbon products. Thecatalytic reaction may involve volatile organics entering into acatalyst (e.g., a zeolite catalyst) where they are converted into, forexample, p-xylene as well as other hydrocarbons such as aromatics andolefins, in addition to carbon monoxide, carbon dioxide, water, andcoke. Inside or upon contact with the catalyst, the pyrolysis productmay undergo a series of dehydration, decarbonylation, decarboxylation,isomerization, oligomerization, and dehydrogenation reactions that leadto aromatics, olefins, CO, CO₂ and water. The catalysts provided forherein may be particularly suited for producing xylenes with arelatively high selectivity to p-xylene in the xylenes of at least about40%, or at least about 45%, or at least about 50%, or at least about55%, or at least about 60%, or at least about 65%, or at least about70%, or at least about 75%, or at least about 80%, or at least about85%, or at least about 90%.

FIG. 1A includes a schematic illustration of an exemplary chemicalprocess design used to perform catalytic pyrolysis, according to theinventive method. The process may comprise a CFP process. Referring toFIG. 1A, feed stream 10 includes a solid hydrocarbonaceous material thatcan be fed to reactor 20. The solid hydrocarbonaceous material maygenerally comprise at least carbon and hydrogen. In certain solidhydrocarbonaceous materials (e.g. wood), carbon may be the most abundantcomponent by mass, while in others (e.g. glucose) oxygen may be moreabundant than carbon. Certain solid hydrocarbonaceous materials may alsocomprise relatively minor proportions of other elements such as nitrogenand sulfur.

The feed streams to the reactor may be free of olefins, or may containolefins in an insignificant amount (e.g., such that olefins make up lessthan about 1 wt %, less than about 0.1 wt %, or less than about 0.01 wt% of the total weight of reactant fed to the reactor). In otherembodiments, however, olefins may be present in one or more reactantfeed streams.

The solid hydrocarbonaceous material feed composition (e.g., in feedstream 10 of FIG. 1A) may comprise a mixture of solid hydrocarbonaceousmaterial and a catalyst. The mixture may comprise, for example, a solidcatalyst and a solid hydrocarbonaceous material. In other embodiments, acatalyst may be provided separately from the solid hydrocarbonaceousmaterial (e.g., by co-feeding the catalyst via an independent catalystinlet). A variety of catalysts may be used. For example, in someinstances, zeolite catalysts with varying molar ratios of silica toalumina, and/or varying pore sizes and/or pore opening sizes, and/orvarying catalytically active metals and/or metal oxides, may be used.

Moisture 12 may optionally be removed from the solid hydrocarbonaceousfeed composition prior to being fed to the reactor, e.g., by an optionaldryer 14. Removal of moisture from the solid hydrocarbonaceous materialfeed stream may be advantageous for several reasons. For example, themoisture in the feed stream may require additional energy input in orderto heat the solid hydrocarbonaceous material to a temperaturesufficiently high to achieve pyrolysis. Variations in the moisturecontent of the solid hydrocarbonaceous feed may lead to difficulties incontrolling the temperature of the reactor. In addition, removal ofmoisture from the solid hydrocarbonaceous feed can reduce or eliminatethe need to process the water during later processing steps.

The solid hydrocarbonaceous feed composition may be dried until thesolid hydrocarbonaceous feed composition comprises less than about 10%,less than about 5%, less than about 2%, or less than about 1% water byweight. Suitable equipment capable of removing water from the feedcomposition is known to those skilled in the art. As an example, thedryer may comprise an oven heated to a particular temperature (e.g., atleast about 80° C., at least about 100° C., at least about 150° C., orhigher) through which the solid hydrocarbonaceous feed composition maycontinuously, semi-continuously, or periodically pass. The dryer maycomprise a vacuum chamber into which the solid hydrocarbonaceous feedcomposition may be processed as a batch. The dryer may combine elevatedtemperatures with vacuum operation. The dryer may be integrallyconnected to the reactor or may be provided as a separate unit from thereactor.

The particle size of the solid hydrocarbonaceous feed composition may bereduced in an optional grinding system 16 prior to passing the solidhydrocarbonaceous feed to the reactor. The average diameter of theground, solid hydrocarbonaceous feed composition exiting the grindingsystem may comprise no more than about 50%, no more than about 25%, nomore than about 10%, no more than about 5%, no more than about 2% of theaverage diameter of the feed composition fed to the grinding system.Large-particle solid hydrocarbonaceous feed material may be more easilytransportable and less messy than small-particle feed material. On theother hand, in some cases it may be advantageous to feed small particlesof solid hydrocarbonaceous material to the reactor. The use of agrinding system allows for the transport of large-particle solidhydrocarbonaceous feed between the source and the process, whileenabling the feed of small particles to the reactor.

Suitable equipment capable of grinding the solid hydrocarbonaceous feedcomposition is known to those skilled in the art. For example, thegrinding system may comprise an industrial mill (e.g., hammer mill, ballmill, etc.), a unit with blades (e.g., chipper, shredder, etc.), or anyother suitable type of grinding system. The grinding system may comprisea cooling system (e.g., an active cooling systems such as a pumped fluidheat exchanger, a passive cooling system such as one including fins,etc.), which may be used to maintain the solid hydrocarbonaceous feedcomposition at relatively low temperatures (e.g., ambient temperature)prior to introducing the solid hydrocarbonaceous feed composition to thereactor. The grinding system may be integrally connected to the reactoror may be provided as a separate unit from the reactor. While thegrinding step is shown following the drying step in FIG. 1A, the orderof these operations may be reversed in some embodiments. In still otherembodiments, the drying and grinding steps may be achieved using anintegrated unit.

Grinding and cooling of the solid hydrocarbonaceous material may beachieved using separate units. Cooling of the solid hydrocarbonaceousmaterial may be desirable, for example, to reduce or prevent unwanteddecomposition of the solid hydrocarbonaceous feed material prior topassing it to the reactor. The solid hydrocarbonaceous material may bepassed to a grinding system to produce a ground solid hydrocarbonaceousmaterial. The ground solid hydrocarbonaceous material may then be passedfrom the grinding system to a cooling system and cooled. The solidhydrocarbonaceous material may be cooled to a temperature lower thanabout 300° C., lower than about 200° C., lower than about 100° C., lowerthan about 75° C., lower than about 50° C., lower than about 35° C., orlower than about 20° C. prior to introducing the solid hydrocarbonaceousmaterial into the reactor. The cooling system may include an activecooling unit (e.g., a heat exchanger) capable of lowering thetemperature of the solid hydrocarbonaceous material. The two or more ofthe drier, grinding system, and cooling system may be combined in asingle unit. The cooling system may be directly integrated with one ormore reactors.

The hydrocarbonaceous material may be transferred to reactor 20. Thereactor may be used, in some instances, to perform catalytic pyrolysisof at least a portion of the first reactant comprising thehydrocarbonaceous material under reaction conditions sufficient toproduce one or more pyrolysis products. In the illustrative embodimentof FIG. 1A, the reactor comprises any suitable reactor known to thoseskilled in the art. For example, in some instances, the reactor maycomprise a continuously stirred tank reactor (CSTR), a batch reactor, asemi-batch reactor, or a fixed bed catalytic reactor, among others. Insome cases, the reactor comprises a fluidized bed reactor, e.g., acirculating fluidized bed reactor. Fluidized bed reactors may, in somecases, provide improved mixing of the catalyst, solid hydrocarbonaceousmaterial during pyrolysis and/or subsequent reactions, which may lead toenhanced control over the reaction products formed. The use of fluidizedbed reactors may also lead to improved heat transfer within the reactor.In addition, improved mixing in a fluidized bed reactor may lead to areduction of the amount of coke adhered to the catalyst, resulting inreduced deactivation of the catalyst in some cases.

The reactor(s) may have any suitable size for performing the processesdescribed herein. For example, the reactor may have a volume betweenabout 0.1-1 L, 1-50 L, 50-100 L, 100-250 L, 250-500 L, 500-1000 L,1000-5000 L, 5000-10,000 L, or 10,000-50,000 L. In some instances, thereactor may have a volume greater than about 1 L, or in other instances,greater than about 10 L, 50 L, 100 L, 250 L, 500 L, 1,000 L, or 10,000L. Reactor volumes greater than about 50,000 L may also be possible. Thereactor may be cylindrical, spherical, or any other suitable shape.

Higher yields of desired product formation, lower yields of cokeformation, and/or more controlled product formation (e.g., higherproduction of p-xylene relative to other products) may be achieved whenparticular combinations of reaction conditions and system components areimplemented in methods and systems described herein. For example,conditions such as the mass normalized space velocity(ies) (e.g., of thesolid hydrocarbonaceous material and/or the fluidization fluid), thetemperature of the reactor and/or solids separator, the reactorpressure, the heating rate of the feed stream(s), the catalyst to solidhydrocarbonaceous material mass ratio, the residence time of thehydrocarbonaceous material in the reactor, the residence time of thereaction products in the solids separator, and/or the catalyst type (aswell as silica to alumina molar ratio and pore mouth opening size) maybe controlled to achieve beneficial results, as described below.

The reactor(s) may be operated at any suitable temperature. In someinstances, it may be desirable to operate the reactor(s) at intermediatetemperatures, compared to temperatures typically used in many previouscatalytic pyrolysis systems. For example, the reactor may be operated attemperatures of between about 400° C. and about 600° C., between about425° C. and about 500° C., or between about 440° C. and about 460° C.Operating the reactor(s) at these intermediate temperatures may allowone to maximize the amount of desirable products. The invention may notbe limited to the use of such intermediate temperatures, however, and inother embodiments, lower and/or higher temperatures can be used.

The reactor(s) may also be operated at any suitable pressure. Thereactor may be operated at a pressure of at least about 100 kPa, or atleast about 200 kPa, or at least about 300 kPa, or at least about 400kPa. The reactor may be operated at a pressure below about 600 kPa, orbelow about 400 kPa, or below about 200 kPa. The reactor may be operatedat a pressure in the range from about 100 to about 600 kPa, or in therange from about 100 to about 400 kPa, or in the range from about 100 toabout 200 kPa. The invention may not be limited to the use of suchpressures, however, and in other embodiments, lower and/or higherpressures may be employed.

It may be advantageous to heat the feed stream(s) at a relatively fastrate as it enters the reactor. High heating rates may be advantageousfor a number of reasons. For instance, high heating rates may enhancethe rate of mass transfer of the reactants from the bulk solidhydrocarbonaceous material to the catalytic reactant sites. This may,for example, facilitate introduction of volatile organic compoundsformed during the pyrolysis of the solid hydrocarbonaceous material intothe catalyst before completely thermally decomposing the solidhydrocarbonaceous material and/or the second reactant into generallyundesired products (e.g., coke). In addition, high heating rates mayreduce the amount of time the reactants are exposed to low temperatures(i.e., temperatures between the temperature of the feed and the desiredreaction temperature). Prolonged exposure of the reactants to lowtemperatures may lead to the formation of undesirable products viaundesirable decomposition and/or reaction pathways. Examples of suitableheating rates for heating the feed stream(s) upon entering the reactormay include, for example, greater than about 50° C./s, greater thanabout 100° C./s, greater than about 200° C./s, greater than about 300°C./s, greater than about 400° C./s, greater than about 500° C./s,greater than about 600° C./s, greater than about 700° C./s, greater thanabout 800° C./s, greater than about 900° C./s, greater than about 1000°C./s, or greater. In some cases, the reactant(s) may be heated at aheating rate of between about 500° C./s and about 1000° C./s. In someembodiments, the heating rate for heating the feed stream(s) uponentering the reactor may be between about 50° C./s and about 1000° C./s,or between about 50° C./s and about 400° C./s. The invention may notlimited to the use of such heating rates, however, and in otherembodiments, lower and/or higher heating rates can be used.

The mass-normalized space velocity of the hydrocarbonaceous material maybe selected to selectively produce a desired array of fluid hydrocarbonproducts. As used herein, the term “mass-normalized space velocity” of acomponent is defined as the mass flow rate of the component into thereactor (e.g., as measured in g/hr) divided by the mass of catalyst inthe reactor (e.g., as measured in g) and has units of inverse time. Forexample, the mass-normalized space velocity of solid hydrocarbonaceousmaterial fed to the reactor may be calculated as the mass flow rate ofthe solid hydrocarbonaceous material into the reactor divided by themass of catalyst in the reactor. The mass-normalized space velocity of acomponent (e.g., the hydrocarbonaceous material) in the reactor may becalculated using different methods depending upon the type of reactorbeing used. For example, in systems employing batch or semi-batchreactors, wherein the solid hydrocarbonaceous material is not fedcontinuously to the reactor, the solid hydrocarbonaceous material doesnot have a mass-normalized space velocity. For systems in which catalystis fed to and/or extracted from the reactor during reaction (e.g.,circulating fluidized bed reactors), the mass-normalized space velocitymay be determined by calculating the average amount of catalyst withinthe volume of the reactor over a period of operation (e.g., steady-stateoperation).

The mass-normalized space velocity of the hydrocarbonaceous material fedto the reactor may be at a mass normalized space velocity of up to about3 hour⁻¹, or up to about 2 hour⁻¹, or up to about 1.5 hour⁻¹, or up toabout 0.9 hour⁻¹, or in the range from about 0.01 hour⁻¹ to about 3hour⁻¹, or in the range from about 0.01 to about 2 hour⁻¹, or in therange from about 0.01 to about 1.5 hour⁻¹, or in the range from about0.01 to about 0.9 hour⁻¹, or in the range from about 0.01 hour⁻¹ toabout 0.5 hour⁻¹, or in the range from about 0.1 hour⁻¹ to about 0.9hour⁻¹, or in the range from about 0.1 hour⁻¹ to about 0.5 hour⁻¹. Theinvention may not be limited to the use of such mass-normalized spacevelocities, however, and in other embodiments, lower and/or highermass-normalized space velocities can be used.

The residence time of a reactant (e.g., the hydrocarbonaceous material)in the reactor (i.e., the reactant residence time) may be at least about1 second, at least about 2 seconds, at least about 5 seconds, at leastabout 7 seconds, at least about 10 seconds, at least about 15 seconds,at least about 20 seconds, at least about 25 seconds, at least about 30seconds, at least about 60 seconds, at least about 120 seconds, at leastabout 240 seconds, or at least about 480 seconds. In some cases, theresidence time of a reactant (e.g., the hydrocarbonaceous material) inthe reactor may be less than about 5 minutes, or from about 1 second andabout 4 minutes, or from about 2 seconds to about 4 minutes, or fromabout 5 seconds to about 4 minutes, or from about 7 seconds to about 4minutes, or from about 10 seconds to about 4 minutes, or from about 12seconds to about 4 minutes, or from about 15 seconds to about 4 minutes,or from about 20 seconds to about 4 minutes, or from about 30 seconds toabout 4 minutes, or from about 60 seconds to about 4 minutes. Previous“fast pyrolysis” studies have, in many cases, employed systems with veryshort reactant residence times (e.g., less than 2 seconds). In somecases, however, the use of relatively longer residence times may allowfor additional chemical reactions to form desirable products. Longresidence times may be achieved by, for example, increasing the volumeof the reactor and/or reducing the volumetric flow rate of thehydrocarbonaceous materials. It should be understood, however, that insome embodiments described herein, the residence time of the reactant(e.g., hydrocarbonaceous material) may be relatively shorter, e.g., lessthan about 2 seconds, or less than about 1 second.

The contact time of the pyrolysis product (e.g., pyrolysis vapor) withthe catalyst in the reactor may be at least about 1 second, at leastabout 2 seconds, at least about 5 seconds, at least about 7 seconds, atleast about 10 seconds, at least about 15 seconds, at least about 20seconds, at least about 25 seconds, at least about 30 seconds, at leastabout 60 seconds, at least about 120 seconds, at least about 240seconds, or at least about 480 seconds. The contact time may be lessthan about 5 minutes, or from about 1 second and about 4 minutes, orfrom about 2 seconds to about 4 minutes, or from about 5 seconds toabout 4 minutes, or from about 7 seconds to about 4 minutes, or fromabout 10 seconds to about 4 minutes, or from about 12 seconds to about 4minutes, or from about 15 seconds to about 4 minutes, or from about 20seconds to about 4 minutes, or from about 30 seconds to about 4 minutes,or from about 60 seconds to about 4 minutes.

In certain cases where fluidized bed reactors are used, the feedmaterial (e.g., a solid hydrocarbonaceous material) in the reactor maybe fluidized by flowing a fluid stream through the reactor. In theexemplary embodiment of FIG. 1A, a fluid stream 44 is used to fluidizethe feed material in reactor 20. Fluid may be supplied to the fluidstream from a fluid source 24 and/or from the product streams of thereactor via a compressor 26. As used herein, the term “fluid” means amaterial generally in a liquid, supercritical, or gaseous state. Fluids,however, may also contain solids such as, for example, suspended orcolloidal particles. In some embodiments, it may be advantageous tocontrol the residence time of the fluidization fluid in the reactor. Theresidence time of the fluidization fluid may be defined as the volume ofthe reactor divided by the volumetric flow rate of the fluidizationfluid. The residence time of the fluidization fluid may be at leastabout 0.1 second, at least about 0.2 second, at least about 0.5 second,at least about 1 second, at least about 2 seconds, at least about 3seconds, at least about 4 seconds, at least about 5 seconds, at leastabout 6 seconds, at least about 8 seconds, at least about 10 seconds, atleast about 12 seconds, at least about 24 seconds, or at least about 48seconds. The residence time of the fluidization fluid may be from about0.1 second to about 48 seconds, from about 0.2 second to about 48seconds, from about 0.5 second to about 480 seconds, from about 1 secondto about 48 seconds, from about 3 seconds to about 48 seconds, fromabout 5 seconds to about 48 seconds, from about 6 seconds to about 48seconds, from about 8 seconds to about 48 seconds, from about 10 secondsto about 48 seconds, from about 12 seconds to about 48 seconds, or fromabout 24 seconds to about 48 seconds.

Suitable fluidization fluids that may be used in this invention include,for example, inert gases (e.g., helium, argon, neon, etc.), hydrogen,nitrogen, carbon monoxide, and carbon dioxide, among others.

As shown in the illustrative embodiment of FIG. 1A, the products (e.g.,fluid hydrocarbon products) formed during the reaction of the reactants(e.g., the solid hydrocarbonaceous material) exit the reactor via aproduct stream 30. In addition to the reaction products, the productstream may, in some cases, comprise unreacted reactant(s), fluidizationfluid, and/or catalyst. In one set of embodiments, the desired reactionproduct(s) (e.g., liquid aromatic hydrocarbons, olefin hydrocarbons,gaseous products, etc.) may be recovered from an effluent stream of thereactor.

As shown in the illustrative embodiment of FIG. 1A, product stream 30may be fed to an optional solids separator 32. The solids separator maybe used, in some cases, to separate the reaction products from catalyst(e.g., at least partially deactivated catalyst) present in the productstream. In addition, the solids separator may be used, in someinstances, to remove coke and/or ash from the catalyst. In someembodiments, the solids separator may comprise optional purge stream 33,which may be used to purge coke, ash, and/or catalyst from the solidsseparator.

The equipment required to achieve the solids separation and/or decokingsteps can be readily designed by one of ordinary skill in the art. Forexample, the solids separator may comprise a vessel comprising a meshmaterial that defines a retaining portion and a permeate portion of thevessel. The mesh may serve to retain the catalyst within the retainingportion while allowing the reaction product to pass to the permeateportion. The catalyst may exit the solids separator through a port onthe retaining side of the mesh while the reaction product may exit aport on the permeate side of the mesh. Other examples of solidsseparators and/or decokers are described in more detail in Kirk-OthmerEncyclopedia of Chemical Technology (Online), Vol. 11, Hoboken, N.J.:Wiley-Interscience, c2001-, pages 700-734; and C. D. Cooper and F. C.Alley. Air Pollution Control, A Design Approach, Second Ed. ProspectHeights, Ill.: Waveland Press, Inc. c1994, pages 127-149, which areincorporated herein by reference.

The solids separator may be operated at any suitable temperature. Insome embodiments, the solids separator may be operated at elevatedtemperatures. For certain reactions, the use of elevated temperatures inthe solids separator can allow for additional reforming and/or reactionof the compounds from the reactor. This may allow for the increasedformation of desirable products. While not wishing to be bound by anytheory, it is believed that elevated temperatures in the solidsseparator may provide enough energy to drive endothermic reformingreactions. The solids separator may be operated at a temperature of, forexample, between about 25° C. and about 200° C., between about 200° C.and about 500° C., between about 500° C. and about 600° C., or betweenabout 600° C. and about 800° C. In some cases, the solids separator maybe operated at temperatures of at least about 500° C., at least about600° C., at least 700° C., at least 800° C., or higher.

It may be beneficial to control the residence time of the catalyst inthe solids separator. The residence time of the catalyst in the solidsseparator may be defined as the volume of the solids separator dividedby the volumetric flow rate of the catalyst through the solidsseparator. In some cases, relatively long residence times of thecatalyst in the solids separator may be desired in order to facilitatethe removal of sufficient amounts of ash, coke, and/or other undesirableproducts from the catalyst. In addition, by employing relatively longresidence times of the catalyst in the solids separator, the pyrolysisproducts may be further reacted to produce desirable products. Theresidence time and temperature in the solids separator may together beselected such that a desired product stream is produced. The residencetime of the catalyst in the solids separator may be at least about 1second, at least about 5 seconds, at least about 7 seconds, at leastabout 10 seconds, at least about 30 seconds, at least about 60 seconds,at least about 120 seconds, at least about 240 seconds, at least about300 seconds, at least about 600 seconds, or at least about 1200 seconds.Methods for controlling the residence time of the catalyst in the solidsseparator are known by those skilled in the art. For example, in somecases, the interior wall of the solids separator may comprise bafflesthat serve to restrict the flow of catalyst through the solids separatorand/or increase the path length of fluid flow in the solids separator.Additionally or alternatively, the residence time of the catalyst in thesolids separator may be controlled by controlling the flow rate of thecatalyst through the solids separator (e.g., by controlling the flowrate of the fluidizing fluid through the reactor).

The solids separator may have any suitable size. For example, the solidsseparator may have a volume between about 0.1-1 L, 1-50 L, 50-100 L,100-250 L, 250-500 L, 500-1000 L, 1000-5000 L, 5000-10,000 L, or10,000-50,000 L. In some instances, the solids separator may have avolume greater than about 1 L, or in other instances, greater than about10 L, 50 L, 100 L, 250 L, 500 L, 1,000 L, or 10,000 L. Solids separatorvolumes greater than 50,000 L are also possible. The solids separatormay be cylindrical, spherical, or any other shape and may be circulatingor non-circulating. In some embodiments, the solids separator maycomprise a vessel or other unit operation similar to that used for oneor more of the reactor(s) used in the process. The flow path for thecatalyst in the solids separator may comprise any suitable geometry. Forexample, the flow path may be substantially straight. In some cases, thesolids separator may comprise a flow channel with a serpentine,meandering, helical, or any other suitable shape. The ratio of thelength of the flow path of the solids separator (or, in certainembodiments, the path length of the catalyst through the solidsseparator) to the average diameter of the solids separator channel maycomprise any suitable ratio. The ratio may be at least about 2:1, atleast 5:1, at least 10:1, at least 50:1, at least 100:1, or greater.

The solids separator may not be required in all embodiments. Forexample, for situations in which catalytic fixed bed reactors areemployed, the catalyst may be retained within the reactor, and thereaction products may exit the reactor substantially free of catalyst,thus negating the need for a separation step.

The separated catalyst may exit the solids separator via stream 34. Aportion of the separated catalyst may be returned to the reactor via areturn pipe, not shown in FIG. 1A. The catalyst exiting the separatormay be at least partially deactivated. The separated catalyst may be fedto a regenerator 36 in which any catalyst that was at least partiallydeactivated may be re-activated. The regenerator may comprise anoptional purge stream 37, which may be used to purge coke, ash, and/orcatalyst from the regenerator. Methods for activating catalyst arewell-known to those skilled in the art, for example, as described inKirk-Othmer Encyclopedia of Chemical Technology (Online), Vol. 5,Hoboken, N.J.: Wiley-Interscience, c2001-, pages 255-322, which areincorporated herein by reference.

A portion of the catalyst may be removed from the reactor through acatalyst exit port (not shown in FIG. 1A). The catalyst removed from thereactor may be partially deactivated and passed via a conduit intoregenerator 36, or into a separate regenerator (not shown in FIG. 1A).Removed catalyst that has been regenerated may be returned to thereactor via stream 47, or may be returned to the reactor separately fromthe fluidization gas via a separate stream (not shown in FIG. 1A).

An oxidizing agent may be fed to the regenerator via a stream 38, e.g.,as shown in FIG. 1A. The oxidizing agent may originate from any sourceincluding, for example, a tank of oxygen, atmospheric air, steam, amongothers. In the regenerator, the catalyst may be re-activated by reactingthe catalyst with the oxidizing agent. The deactivated catalyst maycomprise residual carbon and/or coke, which may be removed via reactionwith the oxidizing agent in the regenerator. The regenerator in FIG. 1Acomprises a vent stream 40 which may include regeneration reactionproducts, residual oxidizing agent, etc.

The regenerator may be of any suitable size mentioned above inconnection with the reactor or the solids separator. In addition, theregenerator may be operated at elevated temperatures in some cases(e.g., at least about 300° C., 400° C., 500° C., 600° C., 700° C., 800°C., or higher). The residence time of the catalyst in the regeneratormay also be controlled using methods known by those skilled in the art,including those outlined above. The mass flow rate of the catalystthrough the regenerator may be coupled to the flow rate(s) in thereactor and/or solids separator in order to preserve the mass balance inthe system.

The regenerated catalyst may exit the regenerator via stream 42. Theregenerated catalyst may be recycled back to the reactor via recyclestream 47. In some cases, catalyst may be lost from the system orremoved intentionally during operation. Additional “makeup” catalyst maybe added to the system via a makeup stream 46. The regenerated andmakeup catalyst may be fed to the reactor with the fluidization fluidvia recycle stream 47. Alternatively, the catalyst and fluidizationfluid may be fed to the reactor via separate streams.

Referring back to solids separator 32 in FIG. 1A, the reaction products(e.g., fluid hydrocarbon products) may exit the solids separator viastream 48. In some cases, a fraction of stream 48 may be purged viapurge stream 60. The contents of the purge stream may be fed to acombustor or a water-gas shift reactor, for example, to recuperateenergy that would otherwise be lost from the system. In some cases, thereaction products in stream 48 may be fed to an optional condenser 50.The condenser may comprise a heat exchanger which condenses at least aportion of the reaction product from a gaseous to a liquid state. Thecondenser may be used to separate the reaction products into gaseous,liquid, and solid fractions. The operation of condensers is well knownto those skilled in the art. Examples of condensers that may be used aredescribed in more detail in Perry's Chemical Engineers' Handbook,Section 11: “Heat Transfer Equipment.” 8th ed. New York: McGraw-Hill,c2008, which is incorporated herein by reference.

The condenser may also make use of pressure change to condense portionsof the product stream. In FIG. 1A, stream 54 may comprise the liquidfraction of the reaction products (e.g., water, aromatic compounds,olefin compounds, etc.), and stream 74 may comprise the gaseous fractionof the reaction products (e.g., CO, CO₂, H₂, etc.). In some embodiments,the gaseous fraction may be fed to a vapor recovery system 70. The vaporrecovery system may be used, for example, to recover any desirablevapors within stream 74 and transport them via stream 72. In addition,stream 76 may be used to transport CO, CO₂, and/or other non-recoverablegases from the vapor recovery system. The optional vapor recovery systemmay be placed in other locations. For example, in some embodiments, avapor recovery system may be positioned downstream of purge stream 54.One skilled in the art can select an appropriate placement for a vaporrecovery system.

Other products (e.g., excess gas) may be transported to optionalcompressor 26 via stream 56, where they may be compressed and used asfluidization gas in the reactor (stream 22) and/or where they may assistin transporting the hydrocarbonaceous material to the reactor (streams58) or may be used to transport catalyst to the reactor (not shown), ormay be used to transport additional non-solid feeds to the reactor. Insome instances, the liquid fraction may be further processed, forexample, to separate the water phase from the organic phase, to separateindividual compounds, etc.

It should be understood that, while the set of embodiments described byFIG. 1A includes a reactor, solids separator, regenerator, condenser,etc., not all embodiments will involve the use of these elements. Forexample, in some embodiments, the feed stream(s) may be fed to acatalytic fixed bed reactor, reacted, and the reaction products may becollected directly from the reactor and cooled without the use of adedicated condenser. In some instances, while a dryer, grinding system,solids separator, regenerator, condenser, and/or compressor may be usedas part of the process, one or more of these elements may compriseseparate units not fluidically and/or integrally connected to thereactor. In other embodiments one or more of the dryer, grinding system,solids separator, regenerator, condenser, and/or compressor may beabsent. In some embodiments, the desired reaction product(s) may berecovered at any point in the production process (e.g., after passagethrough the reactor, after separation, after condensation, etc.).

The process may involve the use of more than one reactor. For instance,multiple reactors may be connected in fluid communication with eachother, for example, to operate in series and/or in parallel, as shown inthe exemplary embodiment of FIG. 1B. The process may comprise providinga solid hydrocarbonaceous material in a first reactor and pyrolyzing,within the first reactor, at least a portion of the solidhydrocarbonaceous material under reaction conditions sufficient toproduce one or more pyrolysis products. A catalyst may be provided tothe first reactor, and at least a portion of the one or more pyrolysisproducts in the first reactor may be catalytically reacted using thecatalyst under reaction conditions sufficient to produce one or morefluid hydrocarbon products. The process may further comprisecatalytically reacting at least a portion of the one or more pyrolysisproducts in a second reactor using a catalyst under reaction conditionssufficient to produce one or more fluid hydrocarbon products. Aftercatalytically reacting at least a portion of the one or more pyrolysisproducts in the second reactor, the process may comprise the step offurther reacting within the second reactor at least a portion of the oneor more fluid hydrocarbon products from the first reactor to produce oneor more other hydrocarbon products.

In FIG. 1B, the reaction product from reactor 20 may be transported to asecond reactor 20′. Those skilled in the art are familiar with the useof multiple-reactor systems for the pyrolysis of organic material toproduce organic products and such systems are known in the art. WhileFIG. 1B illustrates a set of embodiments in which the reactors are influid communication with each other, in some instances, the two reactorsmay not be in fluid communication. For example, a first reactor may beused to produce a first reaction product which may be transported to aseparate facility for reaction in a second reactor. In some instances, acomposition comprising a solid hydrocarbonaceous material (with orwithout a catalyst) may be heated in a first reactor, and at least aportion of the solid hydrocarbonaceous material may be pyrolyzed toproduce a pyrolysis product (and optionally at least partiallydeactivated catalyst). The first pyrolysis product may be in the form ofa liquid and/or a gas. The composition comprising the first pyrolysisproduct may then be heated in a second reactor, which may or may not bein fluid communication with the first reactor. After the heating step inthe second reactor, a second pyrolysis product from the second reactormay be collected. The second pyrolysis product may be in the form of aliquid and/or a gas. In some cases, the composition comprisinghydrocarbonaceous material that is fed into the first reactor maycomprise, for example, a mixture of a solid hydrocarbonaceous materialand a solid catalyst. The first pyrolysis product produced from thefirst reactor may be different in chemical composition, amount, state(e.g., a fluid vs. a gas) than the second pyrolysis product. Forexample, the first pyrolysis product may substantially include a liquid,while the second pyrolysis product may substantially include a gas. Inanother example, the first pyrolysis product may include a fluid product(e.g., a bio-oil, sugar), and the second pyrolysis product may comprisea relatively higher amount of aromatics than the first pyrolysisproduct. In some instances, the first pyrolysis product may include afluid product (e.g., including aromatic compounds), and the secondpyrolysis product may comprise a relatively higher amount of olefinsthan the first pyrolysis product. In yet another example, the firstpyrolysis product may include a fluid product (e.g., a bio-oil, sugar),and the second pyrolysis product may comprise a relatively higher amountof oxygenated aromatic compounds than the first pyrolysis product.

One or more of the reactors in a multiple reactor configuration maycomprise a fluidized bed reactor (e.g., a circulating fluidized bedreactor, a turbulent fluidized bed reactor, etc.) or, in otherinstances, any other type of reactor (e.g., any of the reactorsmentioned above). For example, the first reactor may comprise acirculating fluidized bed reactor or a turbulent fluidized bed reactor,and the second reactor comprises a circulating fluidized bed reactor ora turbulent fluidized bed reactor in fluid communication with the firstreactor. In addition, the multiple reactor configuration may include anyof the additional processing steps and/or equipment mentioned herein(e.g., a solids separator, a regenerator, a condenser, etc.). Thereactors and/or additional processing equipment may be operated usingany of the processing parameters (e.g., temperatures, residence times,etc.) mentioned herein.

Catalyst components useful in the context of this invention can beselected from any catalyst known in the art, or as would be understoodby those skilled in the art made aware of this invention. Functionally,catalysts may be limited only by the capability of any such material topromote and/or effect dehydration, dehydrogenation, isomerization,hydrogen transfer, aromatization, decarbonylation, decarboxylation,aldol condensation and/or any other reaction or process associated withor related to the pyrolysis of a hydrocarbonaceous material. Catalystcomponents can be considered acidic, neutral or basic, as would beunderstood by those skilled in the art.

The catalyst particles described herein may comprise polycrystallinesolids (e.g., polycrystalline particles) in some cases. The catalystparticles may also comprise single crystals, in some embodiments. Incertain cases, the particles may be distinct and separate physicalobjects that are stand-alone. In other cases, the particles may, atleast at certain points in their preparation and/or use, comprise anagglomerate of a plurality of individual particles in intimate contactwith each other.

A catalyst used in embodiments described herein (e.g., in the feedstream, in the reactor, etc.) may be of any suitable size. In somecases, it may be advantageous to use catalysts comprising relativelysmall catalyst particles, which may, as mentioned previously, in certainembodiments, be in the form of larger catalyst objects that may becomprised of a plurality of agglomerated catalyst particles. In someembodiments, for example, the use of small catalyst particles mayincrease the extent to which the hydrocarbonaceous material may contactthe surface sites of the catalyst due to, for example, increasedexternal catalytic surface area and decreased diffusion distancesthrough the catalyst. In some cases, catalyst size and/or catalystparticle size may be chosen based at least in part on, for example, thetype of fluid flow desired and the catalyst lifetime.

In some embodiments, the average diameter (as measured by conventionalsieve analysis) of catalyst objects, which may in certain instances eachcomprise a single catalyst particle or in other instances comprise anagglomerate of a plurality of particles, may be less than about 5 mm,less than about 2 mm, less than about 1 mm, less than about 500 microns,less than about 60 mesh (250 microns), less than about 100 mesh (149microns), less than about 140 mesh (105 microns), less than about 170mesh (88 microns), less than about 200 mesh (74 microns), less thanabout 270 mesh (53 microns), or less than about 400 mesh (37 microns),or smaller.

The catalyst may comprise particles having a maximum cross-sectionaldimension of less than about 5 microns, less than about 1 micron, lessthan about 500 nm, less than about 100 nm, between about 100 nm andabout 5 microns, between about 500 nm and about 5 microns, between about100 nm and about 1 micron, or between about 500 nm and about 1 micron.Catalyst particles having the dimensions within the ranges notedimmediately above may be agglomerated to form discrete catalyst objectshaving dimensions within the ranges noted above. As used here, the“maximum cross-sectional dimension” of a particle refers to the largestdimension between two boundaries of a particle. One of ordinary skill inthe art would be capable of measuring the maximum cross-sectionaldimension of a particle by, for example, analyzing a scanning electronmicrograph (SEM) of a catalyst preparation. In embodiments comprisingagglomerated particles, the particles should be considered separatelywhen determining the maximum cross-sectional dimensions. In such a case,the measurement may be performed by establishing imaginary boundariesbetween each of the agglomerated particles, and measuring the maximumcross-sectional dimension of the hypothetical, individuated particlesthat result from establishing such boundaries. In some embodiments, arelatively large number of the particles within a catalyst may havemaximum cross-sectional dimensions that lie within a given range. Forexample, in some embodiments, at least about 50%, at least about 75%, atleast about 90%, at least about 95%, or at least about 99% of theparticles within a catalyst have maximum cross-sectional dimensions ofless than about 5 microns, less than about 1 micron, less than about 500nm, less than about 100 nm, between about 100 nm and about 5 microns,between about 500 nm and about 5 microns, between about 100 nm and about1 micron, or between about 500 nm and about 1 micron.

A relatively large percentage of the volume of the catalyst can beoccupied by particles with maximum cross-sectional dimensions within aspecific range, in some cases. For example, in some embodiments, atleast about 50%, at least about 75%, at least about 90%, at least about95%, or at least about 99% of the sum of the volumes of all the catalystused is occupied by particles having maximum cross-sectional dimensionsof less than about 5 microns, less than about 1 micron, less than about500 nm, less than about 100 nm, between about 100 nm and about 5microns, between about 500 nm and about 5 microns, between about 100 nmand about 1 micron, or between about 500 nm and about 1 micron.

In some embodiments, the particles within a catalyst may besubstantially the same size. For example, the catalyst may compriseparticles with a distribution of dimensions such that the standarddeviation of the maximum cross-sectional dimensions of the particles isno more than about 50%, no more than about 25%, no more than about 10%,no more than about 5%, no more than about 2%, or no more than about 1%of the average maximum cross-sectional dimensions of the particles.Standard deviation (lower-case sigma) may be calculated as:

$\sigma = \sqrt{\frac{\left. {{\sum\limits_{{i =}\;}^{n}D_{i}} - D_{avg}} \right)^{2}}{n -}}$

wherein D_(i) is the maximum cross-sectional dimension of particle i,D_(avg) is the average of the maximum cross-sectional dimensions of allthe particles, and n is the number of particles within the catalyst. Thepercentage comparisons between the standard deviation and the averagemaximum cross-sectional dimensions of the particles outlined above canbe obtained by dividing the standard deviation by the average andmultiplying by 100%.

Using catalysts including particles within a chosen size distributionindicated above can lead to an increase in the yield and/or selectivityof aromatic compounds produced by the reaction of the hydrocarbonaceousmaterial. For example, in some cases, using catalysts containingparticles with a desired size range (e.g., any of the size distributionsoutlined above) can result in an increase in the amount of aromaticcompounds in the reaction product of at least about 5%, at least about10%, or at least about 20%, relative to an amount of aromatic compoundsthat would be produced using catalysts containing particles with a sizedistribution outside the desired range (e.g., with a large percentage ofparticles larger than 1 micron, larger than 5 microns, etc.).

Alternatively, catalysts may be selected according to pore size (e.g.,mesoporous and pore sizes typically associated with zeolites), e.g.,average pore sizes of less than about 100 Angstroms, less than about 50Angstroms, less than about 20 Angstroms, less than about 10 Angstroms,less than about 5 Angstroms, or smaller. In some embodiments, catalystswith average pore sizes of from about 5 Angstroms to about 100 Angstromsmay be used. In some embodiments, catalysts with average pore sizes ofbetween about 5.5 Angstroms and about 6.5 Angstroms, or between about5.9 Angstroms and about 6.3 Angstroms may be used. In some cases,catalysts with average pore sizes of between about 7 Angstroms and about8 Angstroms, or between about 7.2 Angstroms and about 7.8 Angstroms maybe used.

As used herein, the term “pore size” is used to refer to the smallestcross-sectional diameter of a pore. The smallest cross-sectionaldiameter of a pore may correspond to the smallest cross-sectionaldimension (e.g., a cross-sectional diameter) as measured perpendicularlyto the length of the pore. In some embodiments, a catalyst with an“average pore size” or a “pore size distribution” of X refers to acatalyst in which the average of the smallest cross-sectional diametersof the pores within the catalyst is about X. It should be understoodthat “pore size” or “smallest cross sectional diameter” of a pore asused herein refers to the Norman radii adjusted pore size well known tothose skilled in the art. Determination of Norman radii adjusted poresize is described, for example, in Cook, M.; Conner, W. C., “How big arethe pores of zeolites?” Proceedings of the International ZeoliteConference, 12th, Baltimore, Jul. 5-10, 1998; (1999), 1, pp 409-414,which is incorporated herein by reference in its entirety. As a specificexemplary calculation, the atomic radii for ZSM-5 pores are about5.5-5.6 Angstroms, as measured by x-ray diffraction. In order to adjustfor the repulsive effects between the oxygen atoms in the catalyst, Cookand Conner have shown that the Norman adjusted radii are 0.7 Angstromslarger than the atomic radii (about 6.2-6.3 Angstroms).

One of ordinary skill in the art will understand how to determine thepore size (e.g., minimum pore size, average of minimum pore sizes) in acatalyst. For example, x-ray diffraction (XRD) can be used to determineatomic coordinates. XRD techniques for the determination of pore sizeare described, for example, in Pecharsky, V. K. et al, “Fundamentals ofPowder Diffraction and Structural Characterization of Materials,”Springer Science+Business Media, Inc., New York, 2005, incorporatedherein by reference in its entirety. Other techniques that may be usefulin determining pore sizes (e.g., zeolite pore sizes) include, forexample, helium pycnometry or low pressure argon adsorption techniques.These and other techniques are described in Magee, J. S. et al, “FluidCatalytic Cracking: Science and Technology,” Elsevier PublishingCompany, Jul. 1, 1993, pp. 185-195, which is incorporated herein byreference in its entirety. Pore sizes of mesoporous catalysts may bedetermined using, for example, nitrogen adsorption techniques, asdescribed in Gregg, S. J. at al, “Adsorption, Surface Area andPorosity,” 2nd Ed., Academic Press Inc., New York, 1982 and Rouquerol,F. et al, “Adsorption by powders and porous materials. Principles,Methodology and Applications,” Academic Press Inc., New York, 1998, bothincorporated herein by reference in their entirety. Unless otherwiseindicated, pore sizes referred to herein are those determined by x-raydiffraction corrected as described above to reflect their Norman radiiadjusted pore sizes.

A screening method may be used to select catalysts with appropriate poresizes for the conversion of specific pyrolysis product molecules. Thescreening method may comprise determining the size of pyrolysis productmolecules desired to be catalytically reacted (e.g., the moleculekinetic diameters of the pyrolysis product molecules). One of ordinaryskill in the art can calculate, for example, the kinetic diameter of agiven molecule. The type of catalyst may then be chosen such that thepores of the catalyst (e.g., Norman adjusted minimum radii) aresufficiently large to allow the pyrolysis product molecules to diffuseinto and/or react with the catalyst. In some embodiments, the catalystsare chosen such that their pore sizes are sufficiently small to prevententry and/or reaction of pyrolysis products whose reaction would beundesirable.

The catalyst may be selected from naturally-occurring zeolites,synthetic zeolites and combinations thereof. The catalyst may be aMordenite Framework Inverted (MFI) type zeolite catalyst, such as aZSM-5 zeolite catalyst. Catalysts comprising ZSM-5 that may be used withor without modification are available commercially. The catalysts thatare provided for herein may comprise acid or catalytically active sites.While not wishing to be bound by theory, it is believed that variousacid sites in ZSM-5 and other zeolites are catalytically active forreactions of the hydrocarbonaceous materials including dehydration,decarbonylation, decarboxylation, isomerization, oligomerization and/ordehydrogenation, hence the terms “acid sites” and “catalytically activesites” may be used interchangeably. Other types of useful zeolitecatalysts may include ferrierite, zeolite Y, zeolite beta, modernite,MCM-22, ZSM-23, ZSM-57, SUZ-4, EU-1, ZSM-11, (S)AlP0-31, SSZ-23,mixtures of two or more thereof, and the like.

The catalyst may comprise, in addition to alumina and silica, one ormore additional metals and/or a metal oxides. Suitable metals and/oroxides may include, for example, nickel, platinum, vanadium, palladium,manganese, cobalt, zinc, copper, chromium, gallium, and/or any of theiroxides, among others. The metal and/or metal oxide can be impregnatedinto the catalyst (e.g., in the interstices of the lattice structure ofthe catalyst), in some embodiments. The metal or metal oxide can beadded to the zeolite by any of a number of techniques known to thoseskilled in the art, such as, but not limited to, impregnation, ionexchange, vapor deposition, and the like. The zeolite may comprise smallamounts of structure stabilizing elements such as phosphorus, lanthanum,rare earths, and the like, typically at levels that are less than about1% by weight of the zeolite. The catalyst may be conditioned beforeoperation in the process by a wide range of techniques known to thoseskilled in the art such as, but not limited to, oxidation, calcination,reduction, cyclic oxidation and reduction, steaming, hydrolysis, and thelike. The metal and/or metal oxide may be incorporated into the latticestructure of the catalyst. For example, the metal and/or metal oxide maybe included during the preparation of the catalyst, and the metal and/ormetal oxide may occupy a lattice site of the resulting catalyst (e.g., azeolite catalyst). As another example, the metal and/or metal oxide mayreact or otherwise interact with a zeolite such that the metal and/ormetal oxide displaces an atom within the lattice structure of thezeolite.

In certain embodiments, a Mordenite Framework Inverted (MFI) zeolitecatalyst comprising gallium may be used. For example, agalloaluminosilicate MFI (GaAlMFI) zeolite catalyst may be used. One ofordinary skill in the art would be familiar with GaAlMFI zeolites, whichmay be thought of as aluminosilicate MFI zeolites in which some of theAl atoms have been replaced with Ga atoms. In some instances, thezeolite catalyst may be in the hydrogen form (e.g., H—GaAlMFI). Thegalloaluminosilicate MFI catalyst may be a ZSM-5 zeolite catalyst inwhich some of the aluminum atoms have been replaced with gallium atoms,in some embodiments.

In some instances, the ratio of moles of Si in the galloaluminosilicatezeolite catalyst to the sum of the moles of Ga and Al (i.e., the molarratio expressed as Si:(Ga+Al)) in the galloaluminosilicate zeolitecatalyst may be at least about 15:1, at least about 20:1, at least about25:1, at least about 35:1, at least about 50:1, at least about 75:1, orhigher. In some embodiments, it may be advantageous to employ a catalystwith a ratio of moles of Si in the zeolite to the sum of the moles of Gaand Al of between about 15:1 and about 100:1, from about 15:1 to about75:1, between about 25:1 and about 80:1, or between about 50:1 and about75:1. In some instances, the ratio of moles of Si in thegalloaluminosilicate zeolite catalyst to the moles of Ga in thegalloaluminosilicate zeolite catalyst may be at least about 30:1, atleast about 60:1, at least about 120:1, at least about 200:1, betweenabout 30:1 and about 300:1, between about 30:1 and about 200:1, betweenabout 30:1 and about 120:1, or between about 30:1 and about 75:1. Theratio of the moles of Si in the galloaluminosilicate zeolite catalyst tothe moles of Al in the galloaluminosilicate zeolite catalyst may be atleast about 10:1, at least about 20:1, at least about 30:1, at leastabout 40:1, at least about 50:1, at least about 75:1, between about 10:1and about 100:1, between about 10:1 and about 75:1, between about 10:1and about 50:1, between about 10:1 and about 40:1, or between about 10:1and about 30:1.

In addition, in some cases, properties of the catalysts (e.g., porestructure, type and/or number of acid sites, etc.) may be chosen toselectively produce a desired product.

It may be desirable, in some embodiments, to employ one or morecatalysts to establish a bimodal distribution of pore sizes. In somecases, a single catalyst with a bimodal distribution of pore sizes maybe used (e.g., a single catalyst that contains predominantly 5.9-6.3Angstrom pores and 7-8 Angstrom pores). In other cases, a mixture of twoor more catalysts may be employed to establish the bimodal distribution(e.g., a mixture of two catalysts, each catalyst type including adistinct range of average pore sizes). In some embodiments, one of theone or more catalysts comprises a zeolite catalyst and another of theone or more catalysts comprises a non-zeolite catalyst (e.g., amesoporous catalyst, a metal oxide catalyst, etc.).

For example, in some embodiments at least about 70%, at least about 80%,at least about 90%, at least about 95%, at least about 98%, or at leastabout 99% of the pores of the one or more catalysts (e.g., a zeolitecatalyst, a mesoporous catalyst, etc.) have smallest cross-sectionaldiameters that lie within a first size distribution or a second sizedistribution. In some cases, at least about 2%, at least about 5%, or atleast about 10% of the pores of the one or more catalysts have smallestcross-sectional diameters that lie within the first size distribution;and at least about 2%, at least about 5%, or at least about 10% of thepores of the one or more catalysts have smallest cross-sectionaldiameters that lie within the second size distribution. In some cases,the first and second size distributions are selected from the rangesprovided above. In certain embodiments, the first and second sizedistributions are different from each other and do not overlap. Anexample of a non-overlapping range is 5.9-6.3 Angstroms and 6.9-8.0Angstroms, and an example of an overlapping range is 5.9-6.3 Angstromsand 6.1-6.5 Angstroms. The first and second size distributions may beselected such that the range are not immediately adjacent one another,an example being pore sizes of 5.9-6.3 Angstroms and 6.9-8.0 Angstroms.An example of a range that is immediately adjacent one another is poresizes of 5.9-6.3 Angstroms and 6.3-6.7 Angstroms.

As a specific example, in some embodiments one or more catalysts is usedto provide a bimodal pore size distribution for the simultaneousproduction of aromatic and olefin compounds. That is, one pore sizedistribution may advantageously produce a relatively high amount ofaromatic compounds, and the other pore size distribution mayadvantageously produce a relatively high amount of olefin compounds. Insome embodiments, at least about 70%, at least about 80%, at least about90%, at least about 95%, at least about 98%, or at least about 99% ofthe pores of the one or more catalysts have smallest cross-sectionaldiameters between about 5.9 Angstroms and about 6.3 Angstroms or betweenabout 7 Angstroms and about 8 Angstroms. In addition, at least about 2%,at least about 5%, or at least about 10% of the pores of the one or morecatalysts have smallest cross-sectional diameters between about 5.9Angstroms and about 6.3 Angstroms; and at least about 2%, at least about5%, or at least about 10% of the pores of the one or more catalysts havesmallest cross-sectional diameters between about 7 Angstroms and about 8Angstroms.

In some embodiments, at least about 70%, at least about 80%, at leastabout 90%, at least about 95%, at least about 98%, or at least about 99%of the pores of the one or more catalysts have smallest cross-sectionaldiameters between about 5.9 Angstroms and about 6.3 Angstroms or betweenabout 7 Angstroms and about 200 Angstroms. In addition, at least about2%, at least about 5%, or at least about 10% of the pores of the one ormore catalysts have smallest cross-sectional diameters between about 5.9Angstroms and about 6.3 Angstroms; and at least about 2%, at least about5%, or at least about 10% of the pores of the one or more catalysts havesmallest cross-sectional diameters between about 7 Angstroms and about200 Angstroms.

In some embodiments, at least about 70%, at least about 80%, at leastabout 90%, at least about 95%, at least about 98%, or at least about 99%of the pores of the one or more catalysts have smallest cross-sectionaldiameters that lie within a first distribution and a seconddistribution, wherein the first distribution is between about 5.9Angstroms and about 6.3 Angstroms and the second distribution isdifferent from and does not overlap with the first distribution. In someembodiments, the second pore size distribution may be between about 7Angstroms and about 200 Angstroms, between about 7 Angstroms and about100 Angstroms, between about 7 Angstroms and about 50 Angstroms, orbetween about 100 Angstroms and about 200 Angstroms. In someembodiments, the second catalyst may be mesoporous (e.g., have a poresize distribution of between about 2 nm and about 50 nm).

In some embodiments, the bimodal distribution of pore sizes may bebeneficial in reacting two or more hydrocarbonaceous feed materialcomponents. For example, some embodiments comprise providing a solidhydrocarbonaceous material comprising a first component and a secondcomponent in a reactor, wherein the first and second components aredifferent. Examples of compounds that may be used as first or secondcomponents include any of the hydrocarbonaceous materials describedherein (e.g., sugar cane bagasse, glucose, wood, corn stover, cellulose,hemi-cellulose, lignin, or any others). For example, the first componentmay comprise one of cellulose, hemi-cellulose and lignin, and the secondcomponent comprises one of cellulose, hemicellulose and lignin. Themethod may further comprise providing first and second catalysts in thereactor. In some embodiments, the first catalyst may have a first poresize distribution and the second catalyst may have a second pore sizedistribution, wherein the first and second pore size distributions aredifferent and do not overlap. The first pore size distribution may be,for example, between about 5.9 Angstroms and about 6.3 Angstroms. Thesecond pore size distribution may be, for example, between about 7Angstroms and about 200 Angstroms, between about 7 Angstroms and about100 Angstroms, between about 7 Angstroms and about 50 Angstroms, orbetween about 100 Angstroms and about 200 Angstroms. In some cases, thesecond catalyst may be mesoporous or non-porous.

The first catalyst may be selective for catalytically reacting the firstcomponent or a derivative thereof to produce a fluid hydrocarbonproduct. In addition, the second catalyst may be selective forcatalytically reacting the second component or a derivative thereof toproduce a fluid hydrocarbon product. The method may further comprisepyrolyzing within the reactor at least a portion of thehydrocarbonaceous material under reaction conditions sufficient toproduce one or more pyrolysis products and catalytically reacting atleast a portion of the pyrolysis products with the first and secondcatalysts to produce the one or more hydrocarbon products. In someinstances, at least partially deactivated catalyst may also be used.

In certain embodiments, a method used in combination with embodimentsdescribed herein includes increasing the catalyst to hydrocarbonaceousmaterial mass ratio of a composition to increase production ofidentifiable aromatic compounds. As illustrated herein, representing butone distinction over certain prior catalytic pyrolysis methods, articlesand methods described herein can be used to produce discrete,identifiable aromatic, biofuel compounds selected from but not limitedto benzene, toluene, propylbenzene, ethylbenzene, methylbenzene,methylethylbenzene, trimethylbenzene, xylenes, indanes, naphthalene,methylnaphthelene, dimethylnaphthalene, ethylnaphthalene, hydrindene,methylhydrindene, and dimethylhydrindene and combinations thereof.

In some embodiments, the reaction chemistry of a catalyst may beaffected by adding one or more additional compounds. For example, theaddition of a metal to a catalyst may result in a shift in selectiveformation of specific compounds (e.g., addition of metal toalumina-silicate catalysts may result in the production of more CO). Inaddition, when the fluidization fluid comprises hydrogen, the amount ofcoke formed on the catalyst may be decreased.

The catalyst may comprise both silica and alumina. The silica (SiO₂) andalumina (Al₂O₃) in the catalyst may be present in any suitable molarratio. For example, in some cases, the catalyst in the feed may comprisea silica (SiO₂) to alumina (Al₂O₃) molar ratio of between about 10:1 andabout 50:1, between about 10:1 and about 40:1, or between about 10:1 andabout 20:1, or about 15:1.

In some embodiments, catalyst and hydrocarbonaceous material may bepresent in any suitable ratio. For example, the catalyst andhydrocarbonaceous material may be present in any suitable mass ratio incases where the feed composition (e.g., through one or more feed streamscomprising catalyst and hydrocarbonaceous material or through separatecatalyst and hydrocarbonaceous material feed streams), comprisescatalyst and hydrocarbonaceous material (e.g., circulating fluidized bedreactors). As another example, in cases where the reactor is initiallyloaded with a mixture of catalyst and hydrocarbonaceous material (e.g.,a batch reactor), the catalyst and hydrocarbonaceous material may bepresent in any suitable mass ratio. In some embodiments involvingcirculating fluidized bed reactors, the mass ratio of the catalyst tohydrocarbonaceous material in the feed stream—i.e., in a compositioncomprising a catalyst and a hydrocarbonaceous material provided to areactor—may be at least about 0.5:1, at least about 1:1, at least about2:1, at least about 5:1, at least about 10:1, at least about 15:1, atleast about 20:1, or higher. In some embodiments involving circulatingfluidized bed reactors, the mass ratio of the catalyst tohydrocarbonaceous material in the feed stream may be less than about0.5:1, less than about 1:1, less than about 2:1, less than about 5:1,less than about 10:1, less than about 15:1, or less than about 20:1; orfrom about 0.5:1 to about 20:1, from about 1:1 to about 20:1, or fromabout 5:1 to about 20:1. Employing a relatively high catalyst tohydrocarbonaceous material mass ratio may facilitate introduction of thevolatile organic compounds, formed from the pyrolysis of the feedmaterial, into the catalyst before they thermally decompose to coke. Notwishing to be bound by any theory, this effect may be at least partiallydue to the presence of a stoichiometric excess of catalyst sites withinthe reactor.

In some embodiments, the articles and methods described herein may beconfigured to selectively produce aromatic compounds (e.g., p-xylene) ina single-stage, or alternatively, a multi-stage pyrolysis apparatus. Forexample, in some embodiments, the mass yield of the aromatic compoundsin the fluid hydrocarbon product may be at least about 18 wt %, at leastabout 20 wt %, at least about 25 wt %, at least about 30 wt %, at leastabout 35 wt %, at least about 39 wt %, between about 18 wt % and about40 wt %, between about 18 wt % and about 35 wt %, between about 20 wt %and about 40 wt %, between about 20 wt % and about 35 wt %, betweenabout 25 wt % and about 40 wt %, between about 25 wt % and about 35 wt%, between about 30 wt % and about 40 wt %, or between about 30 wt % andabout 35 wt %. The mass yield of p-xylene may be at least about 1.5% byweight, or at least about 2% by weight, or at least about 2.5% byweight, or at least about 3% by weight.

As used herein, the “mass yield” of aromatic compounds or p-xylene in agiven product is calculated as the total weight of the aromaticcompounds or p-xylene present in the fluid hydrocarbon product dividedby the weight of the solid hydrocarbonaceous material used in formingthe reaction product, multiplied by 100%.

As used herein, the term “aromatic compound” is used to refer to ahydrocarbon compound comprising one or more aromatic groups such as, forexample, single aromatic ring systems (e.g., benzyl, phenyl, etc.) andfused polycyclic aromatic ring systems (e.g. naphthyl,1,2,3,4-tetrahydronaphthyl, etc.). Examples of aromatic compoundsinclude, but are not limited to, benzene, toluene, indane, indene,2-ethyl toluene, 3-ethyl toluene, 4-ethyl toluene, trimethyl benzene(e.g., 1,3,5-trimethyl benzene, 1,2,4-trimethyl benzene, 1,2,3-trimethylbenzene, etc.), ethylbenzene, methylbenzene, propylbenzene, xylenes(e.g., p-xylene, m-xylene, o-xylene, etc.), naphthalene,methyl-naphthalene (e.g., 1-methyl naphthalene, anthracene,9,10-dimethylanthracene, pyrene, phenanthrene, dimethyl-naphthalene(e.g., 1,5-dimethylnaphthalene, 1,6-dimethylnaphthalene,2,5-dimethylnaphthalene, etc.), ethyl-naphthalene, hydrindene,methyl-hydrindene, and dymethyl-hydrindene. Single ring and/or higherring aromatics may be produced in some embodiments. The aromaticcompounds may have carbon numbers from, for example, C₅-C₁₄, C₆-C₈,C₆-C₁₂, C₈-C₁₂, C₁₀-C₁₄.

In some embodiments, aromatic compounds (especially p-xylene) may beselectively produced when the mass-normalized space velocity of thesolid hydrocarbonaceous material fed to the reactor is up to about 3hour⁻¹, or up to about 2 hour⁻¹, or up to about 1.5 hour⁻¹, or up toabout 0.9 hour⁻¹, or in the range from about 0.01 hour⁻¹ to about 3hour⁻¹, or in the range from about 0.01 to about 2 hour⁻¹, or in therange from about 0.01 to about 1.5 hour⁻¹, or in the range from about0.01 to about 0.9 hour⁻¹, or in the range from about 0.01 hour⁻¹ toabout 0.5 hour⁻¹, or in the range from about 0.1 hour⁻¹ to about 0.9hour⁻¹, or in the range from about 0.1 hour⁻¹ to about 0.5 hour⁻¹. Insome instances, aromatic compounds (especially p-xylene) may beselectively produced when the reactor is operated at a temperature ofbetween about 400° C. and about 600° C. (or between about 425° C. andabout 500° C., or between about 440° C. and about 460° C.). In addition,certain heating rates (e.g., at least about 50° C./s, or at least about400° C./s), high catalyst-to-feed mass ratios (e.g., at least about5:1), and/or high silica to alumina molar ratios in the catalyst (e.g.,at least about 30:1) may be used to facilitate selective production ofaromatic compounds (especially p-xylene). Some such and other processconditions may be combined with a particular reactor type, such as afluidized bed reactor (e.g., a circulating fluidized bed reactor), toselectively produce aromatic and/or olefin compounds.

Furthermore, in some embodiments, the catalyst may be chosen tofacilitate selective production of aromatic products (especiallyp-xylene). For example, ZSM-5 may, in some cases, preferentially producerelatively higher amounts of aromatic compounds. In some cases,catalysts that include Bronsted acid sites may facilitate selectiveproduction of aromatic compounds. In addition, catalysts withwell-ordered pore structures may facilitate selective production ofaromatic compounds. For example, in some embodiments, catalysts withaverage pore diameters between about 5.9 Angstroms and about 6.3Angstroms may be particularly useful in producing aromatic compounds. Inaddition, catalysts with average pore diameters between about 7Angstroms and about 8 Angstroms may be useful in producing olefins. Insome embodiments, a combination of one or more of the above processparameters may be employed to facilitate selective production ofaromatic and/or olefin compounds. The ratio of aromatics to olefinsproduced may be, for example, between about 0.1:1 and about 10:1,between about 0.2:1 and about 5:1, between about 0.5:1 and about 2:1,between about 0.1:1 and about 0.5:1, between about 0.5:1 and about 1:1,between about 1:1 and about 5:1, or between about 5:1 and about 10:1.

In some embodiments, the catalyst to hydrocarbonaceous material massratio in the feed is adjusted to produce desirable products and/orfavorable yields. As such, the catalyst to hydrocarbonaceous materialmass ratio may be, for example, at least about 0.5:1, at least about1:1, at least about 2:1, at least about 5:1, at least about 10:1, atleast about 15:1, at least about 20:1, or higher in some embodiments;or, less than about 0.5:1, less than about 1:1, less than about 2:1,less than about 5:1, less than about 10:1, less than about 15:1, or lessthan about 20:1 in other embodiments.

Furthermore, processes described herein may result in lower cokeformation than certain existing methods. For example, in someembodiments, a pyrolysis product can be formed with less than about 30wt %, less than about 25 wt %, less than about 20 wt %, than about 15 wt%, or less than about 10 wt % of the pyrolysis product being coke. Theamount of coke formed is measured as the weight of coke formed in thesystem divided by the weight of hydrocarbonaceous material used informing the pyrolysis product.

The following non-limiting examples are intended to illustrate variousaspects and features of the invention.

Example

A series of zeolite catalysts are prepared and used in CFP processes forconverting furan, 2-methylfuran (2MF), and pinewood, to fluidhydrocarbon products.

Catalysts

Four catalysts identified as ZSM, GaZSM, SD and GaSD are used. ZSM,which is ZSM-5, (Si/Al=15). GaZSM is made using ion exchange, where 1 gof ZSM is refluxed in 100 mL of an aqueous solution of Ga(NO₃)₃ (0.01 M)at 70° C. for 12 h. After ion exchange, the solution is dried at 110° C.to form a dry powder. The dry powder is calcined under air at 550° C. SDis spray-dried HZSM-5. GaSD is made by incipient wetness impregnation ofSD using a Ga(NO₃)₃ solution (0.43 M). The impregnated GaSD is dried at110° C. and calcined under air at 550° C. The Ga content for GaZSM andGaSD is determined by inductively coupled plasma (ICP) analysis.

Chemical liquid deposition (CLD) employing tetraethylorthosilicate(TEOS) is used to modify the catalysts and thereby reduce theirpore-mouth opening sizes. 1 g of catalyst is dispersed in 25 mL ofhexanes. Then 0.15 mL of TEOS is added. The mixture is refluxed at 90°C. for 1 h with stirring. The catalyst is recovered by centrifuge. Thecatalyst is dried at 100° C. for 2 h and calcined at 500° C. for 4 h indry air. The pore mouth modification process is repeated two more times.The pore mouth modified catalysts may be referred to as “silylated”catalysts and are identified below with “*,” for example, ZSM*, GaZSM*and GaSD*. The modification process may be referred to as a TEOS CLDsilyation process.

The catalyst samples are analyzed by temperature programmed desorptionof isopropyl amine (IPA) or 2,4,6-collidine (2,4,6-trimethylpyridine).Before adsorption, the sample is degassed for 2 h at 823° K. Aftercooling the sample to 393° K, it is exposed for 1 h to He that had beensaturated with isopropylamine or 2,4,6-collidine at room temperature byflowing pure He through a bubbler containing the amine. Then the sampleis held at 393° K with He flow for 2 h to remove physisorbed IPA or2,4,6-collidine. The sample is heated to 973° K at 10° K/min. The totalamount of amine desorbed is used to calculate the total number of acidsites, and the amount of amine that desorbs between about 580° K and650° K is used to calculate the number of Bronsted acid sites for eachcatalyst. Due to the size of 2,4,6-collidine, it does not enter ZSM-5pores. Therefore, desorption of 2,4,6-collidine only detects acid siteson the external surface of the catalyst or in the pores near the poremouth openings. The IPA desorption detects acid sites within the zeolitepores as well as the acid sites on the external surface of the catalystor in the pores near the pore mouth openings. The decrease in the2,4,6-collidine adsorption that occurs with the sylilation treatmentshows the decrease in the number of acid sites on the external surfaceand in or near the pore mouth openings. The decrease in these externalsites is believed to be a factor in the production of m- and o-xyleneand in reducing the re-equilibration of p-xylene formed in the pores tom- or o-xylene, and thus improving the selectivity to p-xylene.

TABLE 1 Acid concentrations obtained by 2,4,6-collidine (kineticdiameter = 7.4 Angstroms) and isopropylamine (IPA, kinetic diameter =5.2 Angstroms) temperature programmed desorption analysis. AdsorbentCollidine Collidine Collidine Collidine IPA IPA IPA IPA Catalyst Ga- Ga-ZSM-5 ZSM-5 ZSM-5 ZSM-5 ZSM-5 ZSM-5 GaZSM--5 GaZSM--5 Acids TotalBrønsted Total Brønsted Total Brønsted Total Brønsted Before silylation0.096 0.0548 0.051 0.0311 1.396 0.651 1.015 0.557 After silylation 0.0410.0344 0.032 0.0265 0.706 0.413 0.676 0.335 Reduction (%) 57 37 37 15 4937 33 40

Catalytic Conversion of Furan and 2MF

The catalytic reactions are carried out in a fixed-bed quartz reactor of0.5 inch (1.27 cm) O.D. The catalyst, which is in the form of afixed-bed of particulate solids, is held in the reactor by a quartzfrit. The catalyst bed is calcined at a temperature of 600° C. with airflowing at a rate of 60 mL/min. After calcination the reactor is purgedby helium at 408 mL/min for 5 min. Furan is pumped into the heliumstream using a syringe pump. Prior to the test run, the furan bypassesthe reactor for 30 min. The helium stream containing the furan is thenswitched to go through the reactor. An air bath condenser is used totrap the heavy products. Gas phase products are collected by air bags.All runs are conducted at atmospheric pressure. No pressure drop isdetected across the catalyst bed. After the reaction process iscompleted, the reactor is purged by helium at a flow rate of 408 mL/minfor 45 seconds at the reaction temperature. The effluent is collectedusing air bags. After each reaction is completed, spent catalyst isregenerated at a temperature of 600° C. using air at a flow rate of 60mL/min. CO formed during regeneration is converted to CO₂ by a copperconverter (copper oxide) at a temperature of 240° C. CO₂ is trapped by aCO₂ trap. Coke yield is determined by measuring the weight change of theCO₂ trap. Gas products are identified by GC-MS (Shimadzu-2010) andquantified by GC-FID/TCD (Shimadzu 2014 for gas samples, and HP-7890 forliquid samples). All hydrocarbons in the gas phase products arequantified by the GC-FID. The CO and CO₂ in the gas phase products arequantified by the GC-TCD. The GC-FID is calibrated by C₂-C₆ normalolefins standards (Scott Specialty Gas, 1000 ppm for each olefin),furan, benzene, toluene, xylenes (gas phase standards are prepared forthese aromatics that can vaporize at room temperature), ethylbenzene,styrene, indene, naphthalene, and benzofuran. The sensitivity of ahydrocarbon is assumed to be proportional to the number of carbonmolecules with similar structure (e.g. styrene vs. methylstyrenes;indene vs. methylindenes). The GC-TCD is calibrated by CO and CO₂standards (Airgas, 6% CO₂ and 14% CO, balanced by helium). Less than0.05% carbon or the products are collected in the condenser. A majorityof the products are in either the gas phase or coke deposited on thecatalyst. Carbon balances close with >90% for all runs.

The reaction conditions for the furan conversion are a temperature of550° C., space velocity (WHSV) of 10.2 h⁻¹, and a partial pressure of 6torr. The furan is pumped with a pumping rate 0.58 mL/h, and the carriergas is maintained at 408 mL/min. The amount of catalyst that is loadedinto the reactor is 53 mg. The reaction process for 2MF is the same asthe process for furan, except that 2% propylene (1.986% propylenebalanced by helium) rather than pure helium is used as the carrier gas.The reaction conditions for 2MF conversion are 600° C., WHSV of 5.7 h⁻¹,and partial pressure of 4.9 ton. 2MF is pumped with a pumping rate 0.57mL/min, and the flow of the carrier gas is maintained at 408 mL/min. Theamount of catalyst that is loaded into the reactor is 92 mg.

The results are shown in Table 2.

TABLE 2 Summary of furan and 2MF conversions over ZSM, ZSM*, GaZSM,GaZSM* and GaSD* Feedstock 2MF 2MF 2MF 2MF Furan Furan Furan Furan FuranCarrier gas 2% 2% 2% 2% propylene propylene propylene propylene He He HeHe He Catalyst ZSM ZSM* GaZSM GaZSM* ZSM ZSM* GaZSM GaZSM* GaSD*Temperature (° C.) 600 600 600 600 550 550 550 550 550 Furan/2MF WHSV(h⁻¹) 5.7 5.7 5.7 5.7 10.2 10.2 10.2 10.2 10.2 P_(furan/2MF) (torr) 4.94.9 4.9 4.9 6.0 6.0 6.0 6.0 6.0 Olefins/furans molar 3.09 3.09 3.09 3.09— — — — — ratio Furan/2MF conversion (%) 99 89 98 74 33 32 41 24 24Propylene conversion (%) 31 22 31 21 — — — — — Overall selectivity (%)CO 5.9 6.6 6.0 9.7 13.1 9.3 13.9 9.5 10.6 CO₂ 0.1 0.5 0.7 0.5 1.6 3.42.0 1.6 2.7 Methane 0.0 0.7 1.0 0.4 0.0 0.0 0.0 0.0 0.0 Olefins 27.827.6 18.4 32.0 14.3 17.0 12.0 11.9 16.3 Aromatics 59.6 53.3 68.5 32.437.4 39.0 42.7 47.7 49.3 Coke 6.2 9.4 4.7 13.6 28.5 23.3 23.9 22.8 16.3Oxygenates 0.4 1.7 0.7 11.5 5.1 7.9 5.5 6.5 4.8 p-Xylene 5.1 14.9 6.64.6 1.0 1.9 0.9 1.2 3.7 Aromatic selectivity (%) Benzene 24.4 20.8 34.432.7 18.0 17.7 26.7 24.9 19.2 Toluene 28.6 23.6 34.4 35.0 20.2 19.5 17.615.2 20.8 Xylene 26.9 30.3 16.5 14.7 5.1 5.6 3.5 2.7 8.5Alkylbenzenes^(a) 4.0 10.4 3.5 2.7 1.5 1.7 0.9 0.5 1.6 Styrenes^(b) 2.64.8 5.0 4.3 7.8 7.7 8.6 5.9 7.3 Indenes^(c) 10.2 7.9 3.6 3.5 23.7 19.723.5 10.1 14.6 Naphthalenes^(b) 3.2 2.4 2.7 7.1 23.8 28.2 19.2 40.8 28.0Olefin selectivity (%) Ethylene 61.6 35.6 55.8 49.6 37.6 26.0 43.1 36.133.5 Propylene — — — — 37.8 44.5 43.6 34.0 39.7 C₄ olefins 28.0 36.726.8 30.3 4.5 6.2 4.0 5.0 5.6 Allene 0.1 0.5 0.7 1.0 3.7 4.1 3.1 9.8 5.7C₅ olefins 7.3 16.3 14.2 15.0 10.8 11.2 4.7 11.0 9.4 C₆ olefins 3.0 8.92.0 3.7 5.5 8.0 1.5 4.1 6.1 C₇ olefins 0.0 2.0 0.5 0.3 0.0 0.0 0.0 0.00.0 Xylenes distribution (%) p-Xylene 32 92 58 96 53 89 57 96 87m-Xylene 49 6 34 3 38 9 35 3 10 o-Xylene 19 2 9 1 10 2 8 1 3^(a)Ethylbenzene and trimethylbenzene ^(b)Styrene and methylstyrenes^(c)Indene, methylindenes, and indane ^(d)Naphthalene,methylnaphthalene, and dihydronaphthalene

Figures A and B show overall p-xylene selectivity and xylenesdistribution, obtained from the conversions of 2MF and furan,respectively. In Figure A the overall p-xylene carbon selectivityobtained from ZSM is 5%. This value is increased to 15% by using ZSM*.The silylation dramatically increases the para selectivity from 32% to92%. Similarly, the silylated GaZSM* also shows a significant increaseof para selectivity from 58% (GaZSM) to 96%. However, the overallp-xylene selectivity for GaZSM* is lower than GaZSM. Table 2 shows thatthe conversion of 2MF for GaZSM* (74%) is lower than that for GaZSM(98%). In addition, the 2MF conversion for ZSM* (89%) is also lower thanthat for ZSM (99%). The decrease of activity is believed to be due tosome active sites in the external surface and in the surface nearpore-openings being eliminated by silica deposition. Ga deposited on ZSMmay have the ability to increase overall aromatics selectivity. This isalso shown in Table 2 (2MF+propylene) where the aromatics selectivity is60% for ZSM and 69% for GaZSM. However, deactivation caused bysilylation on GaZSM* may be more severe than with ZSM*, according to 2MFconversion, suggesting that some active Ga species may be located atthese surfaces and may be killed by silylation. These Ga species imposesspace confinement that causes an increase of para selectivity (Figure A,58% for GaZSM and 32% for ZSM). The silylation of GaZSM further imposesmore space confinement and thus, gives a better para selectivity (96%).

For furan conversion (Figure B), increases of overall p-xyleneselectivity and amongst xylene species towards para are observed fromsilylated catalysts. The para selectivity is increased from 53% for ZSMto 89% for ZSM*, and from 57% for GaZSM to 96% for GaZSM*. Silylationalso causes the activity to decrease as shown in Table 1 where furanconversion is lower in silylated catalysts. The increase activity inGaZSM (41%) comparing with ZSM (33%) may be due to Ga species. However,the lowest furan conversion observed on GaZSM* (24%) is again, due tothe active sites that are killed by silica deposition. The significantdecrease in overall xylene selectivity for the 2MF+propylene reactionsis believed to be due to the furan itself not being a good Diels-Alderreaction agent for xylene production. The results for furan conversionusing GaSD* are shown in Table 2. This table shows that the paraselectivity for this catalyst is 87% and suggests that silylation may beused for an FCC catalyst to increase p-xylene yield from biomassconversion in a fluidized-bed reactor. This is shown below where GaSD*is used for pinewood conversion in a bubbled fluidized-bed reactor.Table 3 shows that an increase in para selectivity from 40% for GaSD to72% for GaSD*.

Catalytic Conversion of Pinewood in a Fluidized-Bed Reactor

CFP of pinewood is conducted in a fluidized bed reactor. The fluidizedbed reactor has a two-inch (5.08 cm) diameter, a height of ten inches(25.4 cm), and is made of 316 stainless steel. Inside the reactor, thecatalyst bed is supported by a distributor plate made of stacked 316stainless mesh (300 mesh). Solid pinewood is introduced into the reactorfrom a sealed feed-hopper. Test runs are conducted using GaSD and GaSD*.Prior to the test runs, the pinewood is ground and sieved to a particlesized ranging between 0.25-1 mm. During the reaction, the catalyst isfluidized by helium gas flowing at 800 standard cubic centimeters (sccm)to enable the reactor to operate in the bubbling flow regime. The hopperand feed chamber are continuously purged with helium at 200 sccm tomaintain an inert environment. The total gas flow through the reactor is1000 sccm helium. Both the reactor and the inlet gas stream are heatedto the reaction temperature (550° C.). The reactor is given two hours toreach this temperature before the reaction is started. The effluent gasleaving the reactor flows through a cyclone to remove entrainedparticles. The effluent then flows into 7 condensers in series toseparate liquid and gas phase products. The first 3 condensers areplaced in an ice-water bath with ethanol inside each condenser as asolvent, and the other 4 condensers are surrounded by a dry ice andacetone bath (−55° C.), without any solvent inside the condensers.Uncondensed gas phase products are collected in air bags at 5, 10, 20,and 30 minutes after the biomass first enters the reactor. The reactiontime is 30 minutes. After 30 minutes the reactor is purged using heliumat a flow rate of 1000 sccm for 30 minutes to remove any CFP productsother than coke on the catalyst. Liquid products are extracted from thecondensers using ethanol. The catalyst is regenerated by using air at800 sccm for 3 hours in addition to the 200 sccm helium from the feedchamber purge. During regeneration, the effluent gases pass through acopper converter where CO is converted to CO₂, and the CO₂ is trapped bya CO₂ trap. Gas phase products are analyzed by a GC-FID/TCD (Shimadzu2014). Liquid samples are analyzed by a GC-FID (HP 7890). Coke yield isobtained by analyzing the weight change of the CO₂ trap. The results areshown in Table 3.

TABLE 3 Summary of pinewood conversions over GaSD and GaSD*. Feed StockSWP SWP SWP Catalyst GaSD GaSD* GaSD* T/° C. 550 550 550 WHSV (spacevelocity)/h⁻¹ 0.35 0.39 0.47 Overall Yield (carbon %) Aromatics 23.214.8 13.3 Olefins 8.9 6.3 5.9 Methane 1.5 3.8 3.3 CO₂ 5.4 9.3 8.2 CO17.2 22.9 19.6 Coke 33.3 33.4 30.0 Total 89.4 90.4 80.3 AromaticSelectivity (%) Benzene 22.0 27.6 25.9 Toluene 29.4 35.8 35.3 Xylenes18.5 15.5 16.3 Naphthalene 14.6 8.9 4.4 Ethylbenzene 2.6 2.3 1.7 Styrene2.2 1.1 1.6 Phenol 1.0 1.0 1.7 Benzofuran 1.3 1.8 2.2 Indene 1.2 2.7 3.7Methylnaphthalene 4.6 1.7 3.1 Xylene distribution (%) p-Xylene 40.0 70.772.4 m-Xylene 30.2 21.3 20.1 o-Xylene 29.8 7.97 7.50 Olefin Selectivity(%) Ethylene 46.0 44.4 44.9 Propylene 47.1 46.3 44.3 Butylene 4.31 4.304.96 Butadiene 2.54 5.00 5.79

While the invention has been explained in relation to variousembodiments, it is to be understood that various modifications thereofwill become apparent to those skilled in the art upon reading thespecification. Therefore, it is to be understood that the inventiondisclosed herein includes any such modifications that may fall withinthe scope of the appended claims.

1. A method for producing a fluid hydrocarbon product comprisingp-xylene from a hydrocarbonaceous material, comprising: feeding thehydrocarbonaceous material to a reactor; pyrolyzing within the reactorat least a portion of the hydrocarbonaceous material under reactionconditions sufficient to produce a pyrolysis product; and catalyticallyreacting at least a portion of the pyrolysis product under reactionconditions in the presence of a zeolite catalyst to produce the fluidhydrocarbon product; the zeolite catalyst comprising pores with poremouth openings and catalytic sites on the external surface of thecatalyst, and an effective amount of a treatment layer derived from asilicone compound to reduce the size of the pore mouth openings and torender at least some of the catalytic sites on the external surface ofthe catalyst inaccessible to the pyrolysis product.
 2. The method ofclaim 1 wherein catalytic sites are positioned in the pores near thepore mouth openings, and the treatment layer renders at least some ofthe catalytic sites in the pores near the pore mouth openingsinaccessible to the pyrolysis product.
 3. The method of claim 1 or claim2 wherein the fluid hydrocarbon product comprises xylenes with ap-xylene selectivity in the xylenes of at least about 40%, or at leastabout 45%, or at least about 50%, or at least about 55%, or at leastabout 60%, or at least about 65%, or at least about 70%, or at leastabout 75%, or at least about 80%, or at least about 85%, or at leastabout 90%.
 4. The method of any of the preceding claims wherein at leastabout 15%, or at least about 25%, or at least about 35%, or at leastabout 45%, or at least about 55%, or at least about 65%, or at leastabout 75%, or at least about 85%, or at least about 90%, or at leastabout 95%, or at least about 98%, or at least about 99%, of thecatalytic sites on the external surface of the catalyst are inaccessibleto the pyrolysis product.
 5. The method of claim 2 wherein at leastabout 20%, or at least about 25%, or at least about 30%, or at leastabout 33%, or at least about 35%, or at least about 40%, or at leastabout 45%, or at least about 49%, or at least about 50%, or at leastabout 55%, or at least about 60%, or at least about 65%, or at leastabout 70%, or at least about 75%, or at least about 80%, or at leastabout 85%, or at least about 90%, or at least about 95%, or at leastabout 99% of the catalytic sites in the pores near the pore mouthopenings are inaccessible to the pyrolysis product.
 6. The method of anyof the preceding claims, wherein the zeolite catalyst comprises silicaand alumina, the silica to alumina molar ratio being in the range fromabout 10:1 to about 50:1, or in the range from about 10:1 to about 40:1,or in the range from about 10:1 to about 20:1, or about 15:1.
 7. Themethod of claim 6, wherein the zeolite catalyst further comprisesnickel, platinum, vanadium, palladium, manganese, cobalt, zinc, copper,chromium, gallium, an oxide of one or more thereof, or a mixture of twoor more thereof.
 8. The method of any of the preceding claims whereinthe silicone compound comprises at least one group represented by theformula


9. The method of any of the preceding claims wherein the siliconecompound is represented by the formula:

wherein R₁ and R₂ independently comprise hydrogen, halogen, hydroxyl,alkyl, alkoxyl, halogenated alkyl, aryl, halogenated aryl, aralkyl,halogenated aralkyl, alkaryl or halogenated alkaryl; and n is a numberthat is at least
 2. 10. The method of claim 9 wherein R₁ and/or R₂comprise methyl, ethyl, or phenyl.
 11. The method of claim 9 or claim 10wherein n is a number in the range from about 3 to about
 1000. 12. Themethod of any of the preceding claims wherein the silicone compound hasa number average molecular weight in the range from about 80 to about20,000, or from about 150 to 10,000.
 13. The method of any of thepreceding claims wherein the silicone compound comprisesdimethylsilicone, diethylsilicone, phenylmethylsilicone,methylhydrogensilicone, ethylhydrogen silicone, phenylhydrogen silicone,methylethyl silicone, phenylethyl silicone, diphenyl silicone,methyltrifluoropropyl silicone, ethyltrifluoropropyl silicone,polydimethyl silicone, tetrachloro-phenylmethyl silicone,tetrachlorophenylethyl silicone, tetrachlorophenylhydrogen silicone,tetrachlorophenylphenyl silicone, methylvinyl silicone, hexamethylcyclotrisiloxane, octamethyl cyclotetrasiloxane, hexaphenylcyclotrisiloxane, octaphenyl cyclotetrasiloxane, or a mixture of two ormore thereof.
 14. The method of any of claims 1 to 8 wherein thesilicone compound comprises a tetraorthosilicate.
 15. The method ofclaim 14 wherein the silicone compound comprisestetramethyl-orthosilicate, tetraethylorthosilicate, or a mixturethereof.
 16. The method of any of the preceding claims wherein thereactor comprises a continuously stirred tank reactor, a batch reactor,a semi-batch reactor, a fixed bed reactor or a fluidized bed reactor.17. The method of any of claims 1 to 15 wherein the reactor comprises afluidized bed reactor.
 18. The method of any of the preceding claimswherein the hydrocarbonaceous material comprises a solidhydrocarbonaceous material, a semi-solid hydrocarbonaceous material, aliquid hydrocarbonaceous material, or a mixture of two or more thereof.19. The method of any of the preceding claims, wherein thehydrocarbonaceous material comprises biomass.
 20. The method of any ofthe preceding claims, wherein the hydrocarbonaceous material comprisesplastic waste, recycled plastics, agricultural solid waste, municipalsolid waste, food waste, animal waste, carbohydrates, lignocellulosicmaterials, xylitol, glucose, cellobiose, hemi-cellulose, lignin, sugarcane bagasse, glucose, wood, corn stover, or a mixture of two or morethereof.
 21. The method of any of the preceding claims wherein thehydrocarbonaceous material comprises pyrolysis oil derived from biomass,a carbohydrate derived from biomass, an alcohol derived from biomass, abiomass extract, a pretreated biomass, a digested biomass product, or amixture of two or more thereof.
 22. The method of any of claims 1 to 18wherein the hydrocarbonaceous material comprises furan and/or2-methylfuran.
 23. The method of any of claims 1 to 18 wherein thehydrocarbonaceous material comprises pinewood.
 24. The method of any ofthe preceding claims, wherein the reactor is at a temperature in therange from about 400° C. to about 600° C., or from about 425° C. toabout 500° C., or from about 440° C. to about 460° C.
 25. The method ofany of the preceding claims, wherein the hydrocarbonaceous material isfed to the reactor at a mass normalized space velocity of up to about 3hour⁻¹, or up to about 2 hour⁻¹, or up to about 1.5 hour⁻¹, or up toabout 0.9 hour⁻¹, or in the range from about 0.01 hour⁻¹ to about 3hour⁻¹, or in the range from about 0.01 to about 2 hour⁻¹, or in therange from about 0.01 to about 1.5 hour⁻¹, or in the range from about0.01 to about 0.9 hour⁻¹, or in the range from about 0.01 hour⁻¹ toabout 0.5 hour⁻¹, or in the range from about 0.1 hour⁻¹ to about 0.9hour⁻¹, or in the range from about 0.1 hour⁻¹ to about 0.5 hour⁻¹. 26.The method of any of the preceding claims wherein the fluid hydrocarbonproduct comprises an aromatic compound, the aromatic carbon molar yieldat least about 22%.
 27. The method of any of claims 1 to 26 wherein thefluid hydrocarbon product comprises an olefinic compound, the olefincarbon molar yield being at least about 3%, or at least about 6%, or atleast about 9%, or at least about 12%.
 28. The method of any of thepreceding claims, further comprising the step of recovering the fluidhydrocarbon product.
 29. The method of any of the preceding claims,wherein the fluid hydrocarbon product further comprises aromaticcompounds and/or olefin compounds.
 30. The method of any of thepreceding claims, wherein the fluid hydrocarbon product furthercomprises benzene, toluene, ethylbenzene, methylethylbenzene,trimethylbenzene, o-xylene, m-xylene, indanes naphthalene,methylnaphthelene, dimethylnaphthalene, ethylnaphthalene, hydrindene,methylhydrindene, dimethylhydrindene, or a mixture of two or morethereof.
 31. The method of any of the preceding claims wherein the massyield of p-xylene in the fluid hydrocarbon product is at least about 1.5wt %, or at least about 2 wt %, or at least about 2.5 wt %, or at leastabout 3 wt %.
 32. The method of any of the preceding claims wherein thereactor is operated at a pressure of at least about 100 kPa, or at leastabout 200 kPa, or at least about 300 kPa, or at least about 400 kPa. 33.The method of any of claims 1 to 31 wherein the reactor is operated at apressure below about 600 kPa, or below about 400 kPa, or below about 200kPa.
 34. The method of any of claims 1 to 31 wherein the reactor isoperated at a pressure in the range from about 100 to about 600 kPa, orin the range from about 100 to about 400 kPa, or in the range from about100 to about 200 kPa.
 35. The method of any of the preceding claims,wherein during the catalytically reacting step a dehydration,decarbonylation, decarboxylation, isomerization, oligomerization and/ordehydrogenation reaction is conducted.
 36. The method of any of thepreceding claims wherein the pyrolysis product is formed with less thanabout 30 wt %, or less than about 25 wt %, or less than about 20 wt %,or less than about 15 wt %, or less than about 10% of the pyrolysisproduct being coke.
 37. The method of any of the preceding claims,wherein the pyrolyzing step and the catalytically reacting steps arecarried out in a single vessel.
 38. The method of any of claims 1 to 36,wherein the pyrolyzing step and the catalytically reacting steps arecarried out in separate vessels.
 39. The method of any of the precedingclaims wherein the residence time for the hydrocarbonaceous material inthe reactor is at least about 1 second, at least about 2 seconds, atleast about 5 seconds, at least about 7 seconds, at least about 10seconds, at least about 15 seconds, at least about 20 seconds, at leastabout 25 seconds, at least about 30 seconds, at least about 60 seconds,at least about 120 seconds, at least about 240 seconds, or at leastabout 480 seconds.
 40. The method of any of the preceding claims whereinthe contact time of the pyrolysis product with the catalyst is at leastabout 1 second, at least about 2 seconds, at least about 5 seconds, atleast about 7 seconds, at least about 10 seconds, at least about 15seconds, at least about 20 seconds, at least about 25 seconds, at leastabout 30 seconds, at least about 60 seconds, at least about 120 seconds,at least about 240 seconds, or at least about 480 seconds.